HP_2010_02 by androsov.info

FEBRUARY 2010

HPIMPACT

SPECIALREPORT

TECHNOLOGY

EIA’s energy outlook

CLEAN FUELS

Evaluating the US’ ethanol industry

Heavy-ends upgrading and dieselization key trends in 2010 refining

Calculate turbulent diffusion jet flames Use neural networks on coker heaters

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FEBRUARY 2010 • VOL. 89 NO. 2 www.HydrocarbonProcessing.com

SPECIAL REPORT: CLEAN FUELS

31 35

Viewpoint Key representatives from the energy industry present their insight on how to achieve balanced energy policy, what is the future for alternative fuels, what part will renewable/biofuels play in the transportation fuel mix, and more

Consider new methods for bottom of the barrel processing—Part 1 Advanced methods use molecule management to upgrade heavy ends M. Motaghi, K. Shree and S. Krishnamurthy

Designing vacuum distillation for deep-cut bitumen service To process bitumen-derived residue, success depends on the engineering design concepts and methods applied M. Grande

Catalytic technology: Options for better hydrogen production

Cover SK Energy’s No.2 Fluidized Catalytic Cracker (FCC) Project with a project budget of US $2 billion was executed in less than 31 months from the beginning of front-end engineering through sustained on-specification production for all units. This revamp allowed upgrading atmospheric residue by first desulfurizing it to less than 0.5% sulfur in residue hydrodesulfurizer followed by a residue FCC. Cracked products are further treated and/or processed in the downstream units to produce low-sulfur FCC gasoline, sweet refinery fuel gas, polymer-grade propylene, LPG, alkylate, light-cycle oil and heavy fuel-oil components.

It is not just steam and feed costs that determine processing methods for hydrogen generation K. Schlögl, L. Xu, A. Düker and W. Kaltner

55 63

Dealing with dieselization Several processing options can provide cost-effective ways to maximize diesel make and limit gasoline production M. Stockle and T. Knight

HPIMPACT

Improve usage of regenerated refining catalysts

17 EIA’s energy outlook

Common-sense guidelines detail when it is beneficial to reuse catalysts G. J. Yeh

17 Taking a long, hard look at ethanol

HEAT TRANSFER

Calculating turbulent diffusion jet flames Following these steps may prevent potential damaging hot spots J. P. Meagher

INSTRUMENTATION

Delayed coker heater analysis using an artificial neural network The model was used to study the effects of the many variables that affect coke formation M. Sharma and A. Ponselva

Implement a constrained optimal control in a conventional level controller—Part 2 Novel tuning method enables a conventional PI controller to explicitly handle the three important operational constraints of a liquid level loop in an optimal manner as well as copes with a broad range of level control from tight to averaging control M. Lee, J. Shin and J. Lee

DEPARTMENTS 7 HPIN BRIEF • 15 HPIN ASSOCIATIONS • 17 HPIMPACT • 21 HPIN CONSTRUCTION • 86 HPI MARKETPLACE • 89 ADVERTISER INDEX

COLUMNS 9 HPIN RELIABILITY How pump installation must differ from the way pumps were shipped 11 HPIN EUROPE Forget dieselization of the marine fleet, what about gas? 13 HPINTEGRATION STRATEGIES Control in the field enables high-availability control 25 HPI VIEWPOINT Engineering Earth’s thermostat with CO2? 90 HPIN CONTROL Back to the future: Process control in 2010

Because Hydrocarbon Processing is edited specifically to be of greatest value to people working in this specialized business, subscriptions are restricted to those engaged in the hydrocarbon processing industry, or service and supply company personnel connected thereto.

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Publisher Bill Wageneck bill.wageneck@gulfpub.com EDITORIAL Editor Les A. Kane Senior Process Editor Stephany Romanow Process Editor Tricia Crossey Reliability/Equipment Editor Heinz P. Bloch News Editor Billy Thinnes European Editor Tim Lloyd Wright Contributing Editor Loraine A. Huchler Contributing Editor William M. Goble Contributing Editor Y. Zak Friedman Contributing Editor ARC Advisory Group (various)

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HPIN BRIEF BILLY THINNES, NEWS EDITOR

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In an effort to reduce the need for US taxpayers to fund the cleanup of problematic environmental pollution spots, the Environmental Protection Agency (EPA) has identified additional industry sectors for which it will begin the regulatory development process for any necessary financial assurance requirements. These sectors include the chemical manufacturing industry and the petroleum industry with a specific focus on refineries. Financial assurance requirements are designed to ensure that owners and operators of facilities are able to pay for cleanup of environmental releases and help reduce the number of sites that need to be cleaned up by federal taxpayers through the Superfund program. According to the EPA, it selected these industries based on a variety of information sources, including sites listed on the Superfund National Priorities List and data on hazardous waste generation from the National Biennial Report and data from the Toxics Release Inventory.

Total US petroleum deliveries for December 2009 were up 0.6% from December 2008, according to a recent release from the API. This could reflect an improved economy and possibly increased demand from colder weather. December deliveries, which averaged 19.3 million bpd, outpaced both four-quarter average deliveries of 19 million bpd and full-year average deliveries of 18.7 million bpd. Gasoline deliveries followed a similar pattern, with December deliveries (at an average 9.1 million bpd) rising 2.3% from December 2008. Gasoline deliveries also were up 1.1% for the fourth quarter, and up 0.3% for full-year 2009. “Clearly, petroleum demand is mirroring the economic recovery,” said API Chief Economist John Felmy. “We are seeing December demand figures stronger than fourth quarter figures and fourth quarter figures stronger than full-year figures. But the data also indicates that the recovery still has a distance to go, particularly if you look at ultra-low sulfur diesel,” he said. Mr. Felmy explained that deliveries of ultra-low sulfur diesel were down 11% in December 2009, compared with December 2008. Import levels of crude and products in 2009 lagged prior year levels, with December’s imports of 10.7 million bpd, down 15%, fourth quarter imports of 10.9 million bpd, down 15.4%, and full-year 2009 imports of 11.7 million bpd, down 9.2%. For full-year 2009, crude oil production was up 7% over the prior year, averaging 5.3 million bpd. The US Bureau of Land Management (BLM) is undertaking reforms that follow the recommendations of an interdisciplinary review team which studied a controversial 2008 oil and gas lease sale in Utah. Under the reformed oil and gas leasing policy, BLM will provide: • Comprehensive interdisciplinary reviews that take into account site-specific considerations for individual lease sales. Resource Management Plans will continue to provide programmatic-level guidance, but individual parcels nominated for leasing will undergo increased scrutiny, including internal and external coordination, public participation and interdisciplinary review of available information. • Greater public involvement in developing Master Leasing and Development Plans for areas where intensive new oil and gas extraction is anticipated so that other important natural resource values can be fully considered prior to making an irreversible commitment to develop an area. • Leadership in identifying areas where new oil and gas leasing will occur. The bureau will continue to accept industry expressions of interest regarding where to offer leases, but will emphasize leasing in already-developed areas and will plan carefully for leasing and development in new areas. BLM will also issue guidance regarding the use of categorical exclusions, or CXs, established by the Energy Policy Act of 2005 and that allow the bureau to approve some oil and gas development activities based on existing environmental or planning analysis. Under the new policy, in accordance with White House Council on Environmental Quality guidelines, BLM will not use these CX’s in cases involving “extraordinary circumstances.” HP

■ Method could make refineries more efficient Refineries could trim millions of dollars in energy costs annually by using a new method developed at Purdue University to rearrange the distillation sequence needed to separate crude petroleum into products. According to Rakesh Agrawal, a professor at Purdue, the researchers have demonstrated their method on petroleum plants that separate crude, showing that 70 of the new sequences they identified could enable oil refineries to improve the energy efficiency of this step anywhere from 6% to 48%. “This is important because improving efficiency by 10% at a refinery processing 250,000 bpd would save in excess of $12 million a year if oil were priced at $70 a barrel,” Dr. Agrawal said. “And that’s just a single refinery. For the US petroleum industry as a whole, this is a huge potential savings.” Chemical plants spend from 50% to 70% of their energy in “separations,” which are usually distillation steps required to separate a raw material into various products. In the case of petroleum, four distillation columns are needed to separate raw crude into five separate components—naphtha, kerosene, diesel fuel, gas oil and heavy residue. Distillation is more energy efficient depending on the order in which the columns are operated in the process. The researchers created a computer algorithm that identifies all of the possible sequences and then determines which require the least heat and energy. They used their new technique to determine that there are nearly 6,000 possible sequences for the four columns used in petroleum distillation. Purdue has filed a patent application for the new crude distillation sequences. HP HYDROCARBON PROCESSING FEBRUARY 2010

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HPIN RELIABILITY HEINZ P. BLOCH, RELIABILITY/EQUIPMENT EDITOR HB@HydrocarbonProcessing.com

How pump installation must differ from the way pumps were shipped In general, process pump manufac- ■ Seeing the conveniently baseplate and foundation filled with turers are asked to provide pumps as a epoxy grout. “set” or assembled package comprising mounted-for-shipping package After the epoxy grout has cured, pump, driver and baseplate (Fig. 1). has, over the past few decades, pump and driver are aligned to criteria After ascertaining correct shaft separathat harmonize with best practices— tion to accommodate the selected cou- led to the erroneous assumption essentially the workmanship guidepling, and prealigning the two shaft lines and reliability-focused practices centerlines within perhaps 0.020 in. that the entire package can of modern plants.1,2 (0.5 mm), mounting holes are spotted simply be hoisted up and placed Also, while being aligned, dial indifrom pump and driver to the baseplate’s cators monitor soft-foot and pipe stress; mounting pads. The pump and base- on a suitable foundation. sensitivity to piping being flanged up is plate provider then proceeds to threadclosely monitored. Any dial indicator tap bolt holes that have about a 0.060 in. (1.5 mm) diameter less movement in excess of 0.002 in. will require making corrections to than the mounting holes (“through-holes”) in pump and driver. the piping. The pump is not allowed to become a pipe support. Mounting bolts are inserted at this pump assembly stage and the Finally, it should be noted that best-in-class users specify, and complete pump set is now considered ready for shipment as a pregenerally insist on, epoxy prefilled steel baseplates. Indeed, considaligned “mounted” package to the designated recipient. eration should be given to dispense with the labor-intensive conSeeing the conveniently mounted-for-shipping package has, ventional grouting procedure described above. You can eliminate over the past few decades, led to the erroneous assumption that much of this by purchasing baseplates prefilled with epoxy. They the entire package can simply be hoisted up and placed on a suitrepresent a monolithic block that will never twist and never get able foundation. However, doing so is not best practice; in fact, out of alignment.1 HP how pumps are shipped has relatively little to do with how they LITERATURE CITED should best be installed in the field.1 The old OEM field service 1 Bloch, H. P. and A. R. Budris, Pump User’s Handbook: Life Extension, 2nd personnel knew about this issue, but wise old-field service folks are edition (Fairmont Press, ISBN 0-88173-517-5). no longer employed there. That is why we have to reinvestigate, 2 Bloch, H. P. and F. K. Geitner, Major Process Equipment Maintenance and read and hope that certain erroneous ways we’ve become used to Repair, 2nd edition (Gulf Publishing Company, ISBN 0-88415-663-X). are rediscovered and corrected. Contrary to the understanding of many of today’s pump The author is HP’s Reliability/Equipment editor. A practicing engineer with close manufacturers’ installation or service personnel, best-practices to 50 years of applicable experience, he advises process plants worldwide on relicompanies (BPCs) will not install the equipment as a mounted ability improvement and maintenance cost-avoidance topics. “set.” To ensure level mounting throughout, the baseplate by itself is placed on the foundation into which hold-down bolts or anchor bolts (Fig. 2) were encased when the re-enforced concrete foundations were poured.2 Leveling screws are then used in conjunction with optical laser tools or a machinist’s precision level to bring the baseplate mounting pads into flat and parallel condition side-to-side, end-toend and diagonally, within an accuracy of 0.002 in./ ft (~0.15 mm/m) or better. The nuts engaging the anchor bolts are now being FIG. 1 A baseplate-mounted API-610 secured and the holstyle pump set installed at a low spaces within petrochemical plant (Source: FIG. 2 Steel baseplate with anchor bolt shown on left and Lubrication Systems Company, leveling screw on right. A chock (thick steel washer) is the baseplate as well Houston, Texas). shown between the leveling screw and foundation.1 as the space between HYDROCARBON PROCESSING FEBRUARY 2010

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HPIN EUROPE TIM LLOYD WRIGHT, EUROPEAN EDITOR tim.wright@gulfpub.com

Forget dieselization of the marine fleet, what about gas? Forget smoke-and-mirrors fuel cell engine that exudes only water vapor, a German company has developed one that emits substantial amounts of carbon dioxide (CO2), and this may be a very good thing. I generally deride the automotive fuel cell in this column over hydrogen supply issues. So, what about using a fuel cell based on reformed liquefied natural gas (LNG) at sea to power ships? Environmentally friendly ship. At the recent Copenhagen conference on climate change, there was an R&D iteration of this idea parked next to the Bella Centre. Kristine Bruun, project manager and a senior researcher at DNV Research, told me that the Viking Lady’s fuel cell is used for auxiliary power, but it also provides power to the main engines, albeit a fraction of the total power requirements of this 6,000-ton vessel. But there’s something else about the Viking Lady. Despite being able to run on marine diesel, the vessel’s piston engines have been fueled exclusively with LNG since the ship was delivered in spring 2009. The Wärtsila gas-electric drive means the vessel can take power from both the 8-MW main engines and the 320-kW fuel cell. In this hybrid mode, it’s claimed that the ship can emit 20% less CO2 than a comparable vessel, reduce nitrogen oxide (NOx) emissions by 180 metric tpy— equivalent to 22,000 cars—and, most importantly, practically eliminates sulfur oxide emissions. It’s easy to forget that ships stand at anchor or at berth in harbor towns; here, they become small, fuel oil (FO) or diesel powered floating power stations. There is a greater prize envisioned other than to simply clean up the auxiliary power systems of large ships. The research team foresees an alluring prospect. If it were possible to scale-up the fuel cell so that it provides 100% of the vessel’s energy needs, that would revolutionize efficiency of the vessel. CO2 emissions—carbon ejected from the reformer connected to the fuel cell—would be half those associated with a typical marine drive. Technology advancements. The challenges facing this idea are the immaturity of the fuel-cell technology, marine fleet turnover (retrofitting is a challenge), the space required onboard for a lower energy-content fuel, infrastructure for bunkering, methane boil-off, security of and long-term cost of supply, and, perhaps most critical, LNG availability. An interesting 80/20 effect applies in the fleet turnover question. According to Chris Holmes, a vice president at Purvin and Gertz, a surprisingly small number of vessels account for a large proportion of emissions. According to Mr. Holms, “We conducted a study and found that 70% of the emissions from the marine fleet come from the largest 5% of long-haul vessels. Expand that to 10% of the global long-haul vessel fleet and you’re talking about 85% of the emissions from the marine sector.”

As for fuel storage, “You may be talking about fuel storage four times the size of existing FO tanks,” says Carl Fagergren, who works with strategic issues at Wallenius Marine in Sweden. A large ocean-going container vessel, burning 50 metric tons of FO daily, may have an existing tank capacity of 5,000 m3. So tripling or quadrupling tank capacity implies a clear penalty. But Fagergren doesn’t see the issue as technically insuperable; it is more a matter of economy and infrastructure. Mr. Fagergren says, “There’s a lot of regulation on the way for sulfur and NOx that may well drive the introduction of alternatives such as LNG.” Infrastructure is not necessarily a problem. “We’ve been told that, if the demand is there, the facilities will be ready for us before the ship is built,” says Ms. Bruun. The issues left are methane boil-off, cost and availability. There’s a reason that the FellowShip team is trumpeting CO2 emissions reductions, rather than greenhouse gas (GHG) reductions—methane boil-off. As a much more problematic GHG (CO2 factor is 22), the methane vaporizing in the engines of the Viking Lady is considered by at least one independent consultant, Marintek, to negate up to half of the CO2 emission savings. Even so, Marintek states, “LNG is a very attractive fuel for reducing emissions.” And, says Mr. Holmes, the boil-off issue has been addressed, for example, in the new Q Max and Q Flex vessels. These vessels reliquefy gas and return it to their tanks as boil-off to maintain the extremely cold temperature of LNG. Finally, there is cost and availability. According to Mr. Holmes, “LNG works because people like the Qataris have cheap gas. Remember: Liquefaction consumes energy equivalent to perhaps 10% of the feed gas.” He says it works in Qatar. But costs would prohibit using, say, Russian gas in Europe and liquefying that. Although, he admits that today LNG is a bargain compared to the Viking Lady’s marine-diesel alternative. But, the long-term supply question is problematic. “Today, (LNG) production stands at just under 190 million metric tpy (MM metric tpy) and we expect it to rise in 2012 to just over 260 MM metric tpy,” Holmes says. “But to supply these 10% of ships would seem to require perhaps another 90 MM metric tpy.” If policy makers do bite the bullet on FO, then there may be no better option for shippers in a fuel-constrained world than to move towards LNG. HP Note: For more info on the FellowShip project, visit www. vikinglady.no.

The author is HP’s European Editor. He has been active as a reporter and conference chair in the European downstream industry since 1997, before which he was a feature writer and reporter for the UK broadsheet press and BBC radio. Mr. Wright lives in Sweden and is the founder of a local climate and sustainability initiative.

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HPINTEGRATION STRATEGIES LARRY O’BRIEN, CONTRIBUTING EDITOR lobrien@arcweb.com

Control in the field enables high-availability control Replacing analog 4-20 mA technology with N/A a digital network was a major driver behind 1 Increased plant safety developing Foundation fieldbus technology. 2 3 However, what really differentiates Founda4 5 tion fieldbus from Profibus PA, HART and Reduced cost per control point other process automation technologies, is incorporating a function-block structure and Control accuracy other supporting functions that make Foundation technology a complete infrastructure Single-loop integrity for process automation. A recent ARC white paper outlined this philosophy, describing Foundation fieldbus as a true automation Increased system availability infrastructure that can help HPI and other process plants achieve operational excellence. 0 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 Embedded control functionality in field devices (“field control”), one of the key enablers FIG. 1 On a one-to-five scale, with five being the greatest benefit, users that adopt for achieving high-availability control, repcontrol-in-the-field find the most benefit in single-loop integrity, control accuracy and increased system availability. resents a stepping stone toward single-loop integrity. The premise is simple. With control at the device level, control is truly distributed, and there is truly no billion dollars. Those actually implementing the control-in-thesingle point of failure in the system above the H1 level. If there is a field functionality, however, still represent a small minority. malfunction in the human-machine interface (HMI) and a loss of In ARC’s view, the reason most users don’t actually use controlvisibility into the process, controllers or any other component in in-the-field is because of the perception in the marketplace that it the system and the control loop, including intelligent field devices, compromises reliability and availability, and provides little direct actuators and positioners, and the network, remain unaffected. In economic impact or business value proposition to the manufacturcases where control resides in the distributed control system (DCS), ing enterprise. The reality is far different. field-level control can add another level of redundancy. Many end users have already managed to avoid unplanned downtime when Control-in-the-field business value. Recently, UKfield-level control took over after a failure in the system. based, Industrial Systems and Control Ltd. (ISC, www.isc-ltd. Evidence suggests that control-in-the-field provides an 80% com) released a study called, “Control in the Field: Analysis of increase in mean-time-between-failures (MTBFs) compared to Performance Benefits.” In a series of illustrative simulation studies, traditional DCS control. The increased MTBFs, combined with ISC determined that control-in-the-field has the potential to offer the reduction in data transfers required, substantially increases improved control loop performance due to its ability to offer faster reliability and availability. The overall reduction in network trafsample rates and shorter latencies in the control loop read-executefic also increases network availability (even with the increased write cycle (Fig. 1). ISC examined the differences in timing and amount of device condition, status and other data that must be sequencing associated with control-in-the-field versus a scheme passed to the DCS). employing control in the DCS to establish typical latencies and sample rates that limit control performance. ISC tested many difField-level control provides flexibility. Field-level conferent scenarios and process dynamics and the report outlines the trol means not only increased availability and reliability, but also results and corresponding benefits. While all control-in-the-field increased flexibility. Controllers are free to handle higher-level loops can benefit from increased integrity, flexibility and reliability, control functions such as advanced control and optimization. ISC found that control loop performance benefits can be quite Foundation fieldbus allows for “dynamically instantiable funcsignificant in fast process loops. HP Larry O’Brien is part of the automation consulting team at ARC covering the tion blocks.” This means that users can activate function blocks in process industries, and an HP contributing editor. He is responsible for tracking the different system components as required. In addition to basic PID market for process automation systems (PASs) and has authored the PAS market studcontrol blocks, a large library of other block types is also available. ies for ARC since 1998. Mr. O’Brien has also authored many other market research, These include switches, alarms and so on. strategy and custom research reports on topicsconsulting including process fieldbus, collaborative The author is part of the automation team at ARC covering the partnerships, total automation trends and others. has been with since process industries, and an HPmarket contributing editor. He He is responsible forARC tracking Foundation fieldbus adoption has skyrocketed over the past January 1993,for andprocess started his career withsystems market (PASs) research in the instrumentation the market automation and hasfield authored the PAS several years, particularly in the HPI. ARC estimates that the total markets. market studies for ARC since 1998 market for fieldbus products and services is rapidly approaching a HYDROCARBON PROCESSING FEBRUARY 2010

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Where do You Want to be on the Performance Curve?

P = People M= Methodologies T = Technologies

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HPIN ASSOCIATIONS BILLY THINNES, NEWS EDITOR

bt@HydrocarbonProcessing.com

Association news in brief NPRA files comments opposing ■ The NPRA is EPA’s PSD, greenhouse gas disappointed that the EPA tailoring rule The National Petrochemical & Refiners Association (NPRA) submitted comments in late December to the US Environmental Protection Agency (EPA) opposing its proposed rulemaking on the “prevention of significant deterioration (PSD) and Title V greenhouse gas (GHG) tailoring rule.” The association feels that the EPA’s attempt to regulate GHGs under the Clean Air Act (CAA) “represents the largest expansion of the EPA’s CAA authority since the enactment of the ‘modern’ CAA in 1970. Yet the EPA has failed to provide the public with anything approaching a complete economic analysis of this seminal event.” The NPRA is disappointed that the EPA “has ignored statutory obligations to conduct required and appropriate regulatory analysis, analysis that could outline alternatives to the course of regulation that EPA has chosen or demonstrate less costly approaches to any regulations that may ultimately be required.” The group suggests there are alternative ways to go about dealing with the issue of regulating greenhouse gases. “The EPA could interpret the CAA to trigger only once a National Ambient Air Quality Standard (NAAQS) has been established for a pollutant. Under this interpretation, which is fully consistent with the CAA and EPA’s regulations, the 202 Rule would result in GHGs only being subject to best available control technology requirements if a source otherwise triggers PSD for a criteria pollutant,” the release says. Another suggested approach from the NPRA would have the EPA specify that “under the 202 Rule the date when GHGs are considered subject to ‘actual control’ is when vehicle manufacturers must comply with an attribute-based standard for model year 2012. This will avoid an imminent PSD trigger for stationary sources and give states and EPA more time to address GHG permitting issues.”

“has ignored statutory obligations to conduct required and appropriate regulatory analysis.” InOGE postponed

The International Oil, Gas and Energy Exposition (InOGE) originally scheduled for February 24–25 has been postponed. The event will now take place June 23–24 at the Messe in Berlin. The three stream conference track will bring together executives from leading oil, gas and energy companies and feature open workshops and small, targeted audiences. InOGE believes it is important for operators to have their plants and equipment running at an optimal level. That’s why the expo will have world class maintenance experts talking about how to source the best materials, general maintenance best practices and equipment investments for the future. A recent statement from the conference organizers indicates that they want attendees to be able to: • Learn how their companies can be more cost-efficient, productive and environmentally safe • Adopt a more advanced knowledge of available oil, gas and renewable technological solutions • Engage in networking opportunities to engage with other industry leaders • Expose their companies to domestic and international markets in oil, gas, renewable energy and sustainable technologies • And collaborate with distributors, resellers, financial partners and investors. For more information, visit www. inoge-expo.com.

International Refining Conference debuts in June Hydrocarbon Processing is hosting its inaugural International Refining Conference at the Sheraton Roma Hotel in Italy

June 21–23. This is an event for technical leaders from around the world, offering a high-level technical and operations program. Project engineers, process engineers and management will meet to share knowledge and solutions related to the downstream oil and gas industry, as well as to network with industry peers. Those that are interested in presenting at the conference should send an abstract approximately 250 words in length (including all authors, affiliations, pertinent contact information and proposed speaker) via email to Events@GulfPub.com. Some of the topics to be covered include: deep conversions and heavy-oil technologies; aging equipment and maintenance; advanced catalyst/licensed technology developments; energy efficiency; new carbon and CO2 management technologies; renewables and biofuels; nextgeneration clean fuels; and safety. For more information on the conference and other ways to get involved, please visit www.GulfPub.com/IRC.

ARC’s World Industry Forum is in Orlando Fans of automation, smart plant management and efficiency will be in Orlando, Florida, for ARC’s World Industry Forum taking place February 8–11. In an historic move, ARC is co-locating its three annual forums together under the World Industry Forum moniker. The Operations-focused Forum (which had been in Orlando the previous 13 years), the Asset Lifecycle Management Forum (held in Houston last year) and the Supply Chain and Logistics Forum (held in Boston for many years) are all together in 2010. ARC also encouraged other organizations to plan their events and meetings along with this forum so that attendees can get the most out of their valuable time. Some of the major themes of the Forum will be: “Rethinking Operational Excellence: Innovative Solutions for the Changing Economy;” and “Rethinking Asset Lifecycle Management: Innovative ALM Strategies for Today and Tomorrow.” HP HYDROCARBON PROCESSING FEBRUARY 2010

I 15

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HPIMPACT BILLY THINNES, NEWS EDITOR

BT@HydrocarbonProcessing.com

EIA’s energy outlook The US Energy Information Administration’s (EIA) short-term energy outlook, published in January, offers a variety of interesting tidbits. The EIA projects the price of West Texas Intermediate (WTI)

crude oil will average $80-$84 per barrel in 2010 and 2011. “EIA’s forecast assumes that US real gross domestic product (GDP) grows by 2% in 2010 and by 2.7% in 2011, while world oil-consumption-weighted real GDP grows by 2.5% and 3.7% in 2010 and 2011,” the outlook said.

US diesel fuel and crude oil prices

$5.00

Forecast

Retail diesel Retail heating oil Crude oil

$4.50 $4.00

Dollars per gallon

$3.50 $3.00 $2.50 $2.00 $1.50 $1.00 $0.50 $0.00 Jan 2006

Jan 2007

Jan 2008 Jan 2009 Jan 2010 Retail prices include state and federal taxes

Jan 2011

Source: EIA, Short-Term Energy Outlook, January 2010

FIG. 1

US diesel fuel and crude oil prices from January 2006 to January 2011.

World liquid fuels consumption 5.5

Forecast

Total consumption

Million barrels per day

3.5

Annual growth

2.5

1.5

60 55

0.5

50 45

-0.5

China United States Other countries

-1.5

35 30

-2.5 2003

2004

2005

2006

2007

2008

Source: EIA, Short-Term Energy Outlook, January 2010

FIG. 2

World liquid fuels consumption from 2003 to 2011.

2009

2010

2011

Million barrels per day

4.5

The increase in crude oil prices is expected to drive up the annual average gasoline retail price in the US. Gasoline was $2.35 a gallon in 2009 and the EIA sees that increasing to $2.84 in 2010 and $2.96 in 2011 (Fig. 1). Further, they project gasoline pump prices could exceed $3 per gallon at certain junctures in the upcoming spring and summer months. The EIA sees a global oil demand growth of 1.1 million bpd in 2010 and 1.5 million bpd in 2011 (Fig. 2). It expects non-OECD countries to account for most of the demand growth with a slight demand increase in the US contributing as well. The outlook said that China will continue to lead the world in consumption growth, with a projected increase of 400,000 bpd in both 2010 and 2011. Looking at natural gas consumption in the US, the EIA notes that total consumption fell by 1.5% in 2009. Contrasting with the increases in oil consumption, the EIA believes natural gas will remain flat (Fig. 3). “Higher natural gas prices in 2010 are expected to cause a 2.8% decline in natural gas consumption in the electric power sector in 2010, which will offset growth in the residential, commercial and industrial sectors,” the outlook said. On the subject of CO2 emissions, the EIA sees increases as economic recovery spurs more energy consumption. CO 2 emissions in the US fell by 6.1% in 2009, but will increase 1.5% and 1.7% in 2010 and 2011, respectively. This will be due to increased use of coal in the electric power sector and the expansion of travel-related petroleum consumption. Striking a positive note, though, the outlook notes that even with increased CO2 emissions in 2010 and 2011, the projected emissions in 2011 should still be lower than annual emissions from 1999–2008.

Taking a long, hard look at ethanol The production and use of biofuels is increasing. Ethanol in particular is the lead biofuel in the US, but question have begun to emerge regarding whether ethanol and other biofuels have an overall positive environmental impact, and if the high cost of HYDROCARBON PROCESSING FEBRUARY 2010

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HPIMPACT Natural gas prices 25

20 Dollars per thousand cubic feet

Forecast

Residential price Henry Hub spot price Composite wellhead price

0 Jan 2006

Jan 2007

Jan 2008

Jan 2009

Jan 2010

Jan 2011

Source: EIA, Short-Term Energy Outlook, January 2010

FIG. 3

Natural gas prices in the US from January 2006 to January 2011.

implementing and producing biofuels is really beneficial. The Baker Institute for Public Policy of Rice University recently undertook an analysis of the biofuels program in the US. The report parsed the different government regulations, requirements and subsidies that emerged from the Energy Policy Act of 2005. It also examined logistical and economic challenges to incorporating biofuels into the US’ gasoline pool. Subsidies and tax credits. Those in the biofuels business in the US must be familiar with several federal policies, including the Renewable Fuel Standard, subsidies for blending biofuel with gasoline and tariffs on importing ethanol. The report notes that $3.2 billion in tax credits were given to gasoline blenders in 2007. The US government also doles out a variety of subsidies and economic incentives to gas station owners, some ethanol producers and farmers. Despite this largesse, more than 25 biofuels facilities closed in the first quarter of 2009. Factors contributing to this problem included the economic downturn, the relative high cost of corn for feedstock and excessive transportation expenses. Recognizing this trend, aid for the biofuels industry was included within the economic stimulus package passed by the US Congress last year. Of particular note was the $480 million for pilot and demonstration-scale biorefineries and an additional $176.5 million for existing commercial scale biorefinery projects.

Transportation issues. Most ethanol

production in the US takes place in the Midwest, which has created difficulty in transporting it to the rest of the country. A majority of US states have not achieved a 10% average ethanol content level within gasoline (E-10 fuel). This is because an effective ethanol distribution network has yet to be developed. Ethanol cannot travel in dedicated petroleum pipelines due to the water solubility of ethanol. If one looks to using trucks, rail or barges to transport ethanol, then environmental problems emerge. Moving ethanol by these methods means that any environmental benefits created by using ethanol are offset by the use of oil-based fuel to transport it. Ethanol vs. gasoline. The report

underlines the importance of ethanol costing less than gasoline: “The relative price of ethanol to gasoline is important because it determines the competitiveness of the two fuels. Given that ethanol has a lower heating value than gasoline—hence yielding lower fuel efficiency—ethanol’s price must be no more than roughly two-thirds of the price of gasoline to make it competitive in the marketplace to sell a blended mixture. Only if ethanol is cheaper than gasoline will blenders make a profit by adding ethanol to their fuel. Without government subsidies, the average ethanol price compared to gasoline will not be commercially competitive in most regional markets in the US to incentivize blenders to add ethanol to gasoline.”

One solution to this problem that the report suggests is to remove the tariffs on importing biofuels from other countries. The sugarcane-based biofuel from Brazil has relatively low production costs, at approximately 30% the average production costs when using corn. “This means that imports of sugarcane-based ethanol have a competitive advantage in certain US costal markets,” the report said. The conundrum in this solution is that if tariffs were dropped on Brazilian sugarcane-based ethanol, the price of corn would have to drop below $3 a bushel for the US domestic corn-based ethanol to compete. US refining industry. Oil refiners in the US are exploring using source materials other than corn for ethanol. One example cited in the report was that of ExxonMobil’s joint venture with Synthetic Genomics to developed biofuels from photosynthetic algae. The project is optimistic that algae could yield more than 2,000 gpy of fuel per acre of production. Also noted was Chevron’s work on converting agricultural waste and non-food crops into biofuels. Environmental concerns. Meeting the US government’s mandated 15 billion gallons of ethanol from corn will require 1.7 trillion gallons of irrigation water, the report said. This is 3% of all irrigation water used in the US in 2000. Thus, there is a concern about increased water usage to create this fuel and a subsequent problem that could emerge: the runoff of ethanol and ethanol tainted water into water supplies and an expansion of the problem of leaking underground storage tanks. “The greater risk to human health comes from the potential for benzene, toluene, ethylbenzene and xylenes (BTEX) mixed with ethanol to travel farther and be more difficult to degrade,” the report said. “Of the BTEX hydrocarbons, benzene is potentially the most toxic and is known to be carcinogenic.” This is a concern because the degradation of benzene is inhibited when it is in the presence of ethanol. Future recommendations. The report

recommends that ethanol blender credits should not be extended by Congress (they expired at the end of 2009). It also encourages Congress to “refrain from giving preferential treatment to corn-based ethanol on the basis of its purported ability to reduce greenhouse gas emissions.” Lifting tariffs on ethanol from Latin America is also recommended. HP HYDROCARBON PROCESSING FEBRUARY 2010

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HPIN CONSTRUCTION BILLY THINNES, NEWS EDITOR BT@HydrocarbonProcessing.com

North America Rentech, Inc. and ClearFuels Technology Inc. have been selected to receive up to $23 million as a grant from the US Department of Energy to construct a biomass gasifier at Rentech’s Energy Technology Center in Denver, Colorado. The gasifier will be integrated with Rentech’s product demonstration unit for the production of renewable synthetic fuels from biomass. The 20 tpd biomass gasifier will produce syngas from sugarcane bagasse, virgin wood waste and other cellulosic feedstocks. ZeaChem Inc. has received a $25 million grant from the US Department of Energy. The grant will be used in conjunction with ZeaChem’s existing investment to support construction of the company’s first cellulosic biorefinery, which will have a capacity of 250,000 gpy. The biorefinery is being built in Boardman, Oregon. ZeaChem recently announced that construction is underway on the biorefinery. Jacobs Engineering Group Inc. has a contract from Sasol to provide basic engineering services for the first commercial installation of Sasol’s proprietary tetramerization technology. The initial commercial unit is expected to have a combined capacity for 1-octene and 1-hexene of 100,000 tpy. The location for the first unit has not been determined, however, Sasol’s chemical complex, located near Lake Charles, Louisiana, is a leading candidate. The basic engineering phase is expected to be completed by the third quarter of 2010. Construction is expected to begin in 2011 with initial production slated for 2013. Enerkem Corp, has been selected by the US Department of Energy to receive $50 million in funding for the construction and operation of a waste-to-biofuels facility to be located in Pontotoc, Mississippi. UOP LLC has been selected for a $25 million award from the US Department of Energy to build a demonstration unit in Hawaii to convert cellulosic biomass into transportation fuels. This project will use cellulosic biomass feedstocks like forestry and agricultural residuals or algae residu-

als to produce the liquid biofuel pyrolysis oil, that will then be upgraded into green transportation fuels. The demonstration plant, which will be built at the Tesoro Corp. refinery in Kapolei, Hawaii, is expected to start up in 2014. The demonstration unit will employ the rapid thermal processing (RTP) technology developed by Ensyn Corp. RTP rapidly heats biomass at ambient pressure to generate high yields of pourable, liquid pyrolysis oil. The pyrolysis oil will then be upgraded to transport fuels using technology developed by UOP. KBR has a contract with International Alliance Group on behalf of Consumers’ Co-operative Refineries Ltd. (CCRL) to provide fabricated pipe spools for CCRL’s industrial facility in Regina, Saskatchewan, Canada. KBR will fabricate pipe spools in a controlled manufacturing environment, with the work being executed in KBR’s Canadian fabrication facility in Edmonton, Alberta, Canada.

TREND ANALYSIS FORECASTING Hydrocarbon Processing maintains an extensive database of historical HPI project information. Current project activity is published three times a year in the HPI Construction Boxscore. When a project is completed, it is removed from current listings and retained in a database. The database is a 35-year compilation of projects by type, operating company, licensor, engineering/constructor, location, etc. Many companies use the historical data for trending or sales forecasting. The historical information is available in comma-delimited or Excel® and can be custom sorted to suit your needs. The cost of the sort depends on the size and complexity of the sort you request and whether a customized program must be written. You can focus on a narrow request such as the history of a particular type of project or you can obtain the entire 35-year Boxscore database, or portions thereof. Simply send a clear description of the data you need and you will receive a prompt cost quotation. Contact: Lee Nichols P. O. Box 2608 Houston, Texas, 77252-2608 Fax: 713-525-4626 e-mail: Lee.Nichols@gulfpub.com.

South America Petrobras has awarded several contracts for the construction of the Abreu e Lima refinery in the state of Pernambuco, Brazil. The contracts together total R$8.9 billion. For the construction of delayed coker units, Petrobras allocated R$3.4 billion to Camargo Correa/CNEC. Conest-UHDT is slated to implement diesel and naphtha hydrotreatment units and hydrogen generation units for R$3.19 billion. Odebrecht Plantas Industriais e Participacoes SA and Construtora OAS Ltda. are on tap to deploy atmospheric distillation units for R$1.48 billion. Conduto-Egesa earned the contract for the pipelines to the refinery, while Construcap-Progen will be providing civil infrastructure services at a price tag of R$120 million. When complete, the refinery is expected to process 230,000 bpd of heavy oil.

Europe ABB has an order worth $26 million from Hellenic Petroleum SA to provide an integrated power and automation system for the upgrade of Hellenic Petroleum’s Elefsina refinery in Greece. ABB will design, supply, install and commission the electrical and automation system to power the refinery. The turnkey electrical solution aims to strengthen the reliability and quality of power supply to the refinery, while improving energy efficiency and reducing overall electricity consumption and costs. The project is expected to be completed in 2010. Foster Wheeler AG’s Global Power Group has a contract to design, supply and erect a heat recovery steam generator (HRSG) for Repsol Petroleo SA. The HRSG and auxiliary equipment will be integrated in a cogeneration plant being built at the Repsol Cartagena refinery in Murcia, Spain. Foster Wheeler will also provide startup supervision services. Galp Energia and Tecnicas Reunidas have agreed on a lump-sum turnkey contract for the conversion of the Sines refinery in Portugal. The contract is worth €1.08 million and includes the conversion project and undertakings in energy efficiency, reliability and environmental areas. Completion is expected in the third quarter of 2011. HYDROCARBON PROCESSING FEBRUARY 2010

I 21

HPIN CONSTRUCTION The goal of the refinery’s conversion project is to maximize the production of diesel from 2011 into the future. The project includes the construction of a new hydrocracker for diesel and aviation fuel, along with a new steam reformer for producing hydrogen and a unit for recovering sulfur.

ment (EPCM) services to upgrade an underground gas storage facility in Cere-la-Ronde, France. This renovation project is expected to increase storage capacity at the Cere-laRonde site up to 1,200 million cubic meters as well as reduce its environmental impact.

Middle East Jacobs Engineering Group Inc. has a contract with Storengy to provide engineering, procurement and construction manage-

Siemens Energy has an order from Samsung Engineering Co. Ltd. for five compressor trains, that will be employed

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Select 152 at www.HydrocarbonProcessing.com/RS 22

Middle East

Graham Corp. has been awarded two orders from refinery customers totaling approximately $3 million. One order is for an upgrade to an existing Graham ejector system at a US refinery that is being reengineered to expand the refiner’s capability to process a wider variety of crude feedstock. The order is expected to ship in the spring of 2010. The second order is for customengineered steam surface condensers to be installed at a large oil refinery currently under construction in the Middle East. This is Graham’s third order related to the project. The two previous orders are currently in backlog. Shipment of the condensers is scheduled for the fourth quarter of FY2011. Alfa Laval has an order for heat exchangers from an Algerian refinery. The order value is about SEK 60 million and was booked in late December 2009. Delivery is scheduled for 2011.

Asia-Pacific

Mining and Metallurgy

North America

in an air separation plant operated by Air Products in Al Jubail, Saudi Arabia. The end customer is the National Industrial Gases Co. Delivery of the first machines is scheduled for the fall of 2010.

Oceania

Chiyoda Singapore (Pte) Ltd. has a contract with Shell Eastern Petroleum Ltd. for a basic design and engineering package for desulfurization facilities at the Pulau Bukom refinery in Singapore. The work will be executed by Chiyoda Singapore (Pte) Ltd. as prime contractor, with the collaboration of Chiyoda Corp. The contract on a unit rate basis is for nine months and comes with an option for an additional three years. Alfa Laval has an order to provide plate heat exchangers for an integrated refinery and petrochemical complex in China. The order value is about SEK 90 million. The plate heat exchangers will be used in the production of mixed xylene. Petrochina Jilin Petrochemical Co. (Jilin) has awarded a contract to Lummus Technology for the license and process design of a grassroots ethylbenzene and styrene monomer (EB/SM) plant in Jilin, China. It has a design capacity of 320,000 metric tpy of styrene monomer and will utilize proprietary technologies provided by Lummus/UOP to minimize costs. The plant, expected to start up in 2011, is the second EB/SM plant awarded to Lummus Technology by Jilin. HP

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Engineering Earth’s thermostat with CO2? Earth’s temperature is a chemical process system. Review of control system engineering of Earth’s thermostat with anthropogenic CO2 in 1997 proved it will never work because it is an unmeasurable, unobservable and uncontrollable system. CO2 does not affect temperature; temperature affects CO2. There are no greenhouse gases in physics. CO2 is not a pollutant; it is green plant food. Global warming stabilized since 1998. Purpose. Physics, chemistry, mathematics, engineering, eco-

nomics, history and ethics are deployed to identify the barriers to designing the thermostat to control Earth’s atmospheric temperature by adjusting its CO2 input. Method. People have beliefs and knowledge. Knowledge of

nature is discovered by the scientific method: theory in the language of nature (mathematics), prediction and verification. Such discoveries are held to be true until falsified. I offer claims supported by credible evidence, settled science and warrant how one can know that they are true. Sound engineering requires no less. No opinion, no speculation, no controversy, no politics, no alarmist adjectives. When an unlicensed engineer resorts to name calling and threats, I declare victory and move on. Credentials. I built a thermostat to verify my “Time Optimum

Control of Chemical Processes” PhD Thesis theory at Purdue in 1966; the first computer control loop in Shell Oil Co., a FCC regenerator thermostat at Deer Park, Texas in 1967; and digital autopilots and spacecraft trajectory controls for NASA’s Apollo Program in 1968. I invented and commercialized hundreds of true boiling cut-point thermostats for petroleum product quality in the hydrocarbon processing industry (HPI) since 1970. I am a registered PE chemical engineer in Texas and control system engineer in California. I was Control Engineer of the Year 1999 and Purdue’s Outstanding Chemical Engineer 2007. I am a contributor to the US Senate Minority Report, “700 Scientists Dissent and Debunk Man-Made Global Warming,” March 16, 2009. I personally financed this presentation; I have no financial incentive in the outcome. I seek no government or business funding. I am an anthropogenic global warming (AGW) skeptic denier. Science.

• CO2 is not a pollutant; it is harmless green plant food. CO2 is the inert result of complete oxidation. There are only two CO2 gas phase reactions, both are endothermic: arc welding and photosynthesis (CO2 + H2O + sunlight = sugars + O2, catalyzed by chlorophyll). US Navy submarines limit CO2 to < 8,000 ppmv because it displaces O2. • Halting all combustion of hydrocarbons (oil, gas, coal and wood) by man will not measurably affect atmospheric CO2 content, now 380 ppm. A simple material balance shows man gener-

ates 30 billion tons/year (this is neither a big nor a small number, it is just a number) while plants consume 7 trillion tons/year (this is neither a big nor a small number, it is just a number). Forest fires, rotting flora and volcanoes input most of the CO2 to the atmosphere. Total input or output is >7. The ratio is 0.03/7 = 0.0043 (this is a small ratio). Cutting the 30 in half to 15 will drop CO2 by 100 ppm after 70 years. • CO2 does not affect temperature; rather temperature affects CO2. Data for the past 400,000 years, reported by Al Gore, An Inconvenient Truth in 2005, shows they cycle together but CO2 lags temperature by about 800 years. Solubility of CO2 in water, oceans, beer and champagne decreases with temperature so solar warming of the ocean releases dissolved CO2 and cooling reabsorbs it. Solar radiation drives Earth’s temperature; CO2 has nothing to do with it. • Atmospheric radiation absorption and emission are dominated by the presence of all three phases of H2O. Like all molecules, CO2 only absorbs and emits specific spectral wavelengths (14.77 microns) that constitute a tiny fraction of solar radiation energy in Earth’s atmosphere. The first 50 ppm of CO2 absorbs about half of this tiny energy, each additional 50 ppm absorbs half of the remaining tiny fraction, so at the current 380 ppm there are almost no absorbable photons left. CO2 could triple to 1,000 ppm with no additional discernable absorption–emission. This is the Beer-Lambert Law: The intensity of radiation decreases exponentially as it passes through an absorbing medium. • There is no such thing as a greenhouse gas because the atmosphere has no glass house. German physicists Gerhard Gerlich and Ralf D Tscheuschner proved this in their classic paper, “Falsification of The Atmospheric CO2 Greenhouse Effects Within The Frame of Physics,” International Journal of Modern Physics B, v23, n03, January 6, 2009, pp. 275-364. Free download at http://arxiv. org/PS_cache/arxiv/pdf/0707/0707.1161v4.pdf. • Earth’s temperature increased naturally 0.6°C from 1976 to 1998 and has stabilized since, decreasing nearly 0.1°C from 2005 to 2009. Forecasts of long-term cooling are credible but irrelevant to the claim anthropogenic CO2 does not affect temperature. CO2 content did not accelerate at the onset of the increase in hydrocarbon combustion by man after 1900. • Warming or cooling, the surface temperature rate of change at a moment in time, does not affect the melting or freezing rate of H2O, only the average temperature of its surroundings > 0°C or < 0°C does. If average temperature is < 0°C, water will freeze even if the temperature is increasing; if average temperature is > 0°C, ice will melt even if the temperature is decreasing. In other words Tim Lloyd Wright is HP’s European Editor and has been active as a reporter Theconference author, president of CLIFFTENT is an independent consulting and chair in the European Inc., downstream industry since 1997,chemical before engineer in identifying, capturing measurable financial which he specializing was a feature writer and reporter forand thesustaining UK broadsheet press and BBC value Mr. fromWright HPI dynamic process and control, IT andofCIM solutions using radio. lives in Sweden is founder a local climate(CLIFFTENT) and sustainability performance-based shared risk–shared reward (SR2) technology licensing. initiative.

HYDROCARBON PROCESSING FEBRUARY 2010

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VIEWPOINT ice melts because its surroundings are too warm, not because they are warming. This is calculus, Isaac Newton, Principia, 1687. • Earth’s atmospheric temperature is not measurable. Temperature is a point property of the energy content of vibrating and radiating molecules. Physics has no rigorous definition of average temperature of bulk matter, accounting for changes in temperature, state, composition, pressure, heat capacity, velocity and reactions. Air temperature and pressure change with altitude, latitude, clouds, time of day, season, weather fronts and deforestation. Wien’s Law gives an average surface temperature from radiation emitted by black bodies like stars; it does not apply to bodies dominated by nonuniform, variable reflection, like Earth. • UN IPCC climate models incorrectly assume Earth’s radiation to space decreases as its temperature increases. The StefanBoltzmann Law states all bodies radiate proportional to T**4. In July 2009, Prof. Richard Lindzen, MIT meteorologist, verified Earth obeys this law. Control engineers know all matter reaches an equilibrium temperature due to this change-mitigating effect. Otherwise Earth would have exploded or frozen long ago. UN climate models are empirical and hence, wrong. • The Waxman-Markey HR2454 CO2 cap-and-trade bill requires the USA to reduce its CO2 production by 83% from 2010 to 2050. Using discredited empirical UN IPCC models, they predict this will reduce CO2 by 20 ppm and temperature by 0.05°C after 40 years. Physics predicts the temperature change approaches zero. • Sea level is changing slowly and naturally in direct proportion to land ice changes, not floating sea ice. Archimedes proved his buoyancy law about 250 BC. Most Northern Hemisphere glaciers have been receding since the ice age ended 18,000 years ago. They have not accelerated since 1900. All AGW scares, like hurricanes, droughts and dying polar bears, have been competently debunked. • Arctic ice shrinks annually when Earth is too warm, but Siberian and Canadian snowfall increases, increasing Northern Hemisphere solar reflectivity, causing Earth to cool and ice to grow again. A plausible mechanism for these regular 40,000-year ice age cycles has been related to the shallowness of the Barents Sea south of Spitsbergen where the Gulf Stream can break through to the Arctic Ocean. Data indicate another regular ice age began since 2000. CO2 is not involved. Engineering.

• Earth’s temperature system cannot be adequately modeled for control. Modeling and control of multivariable, nonlinear, dynamic systems like fluid catalytic cracking, crude distillation, coking, hydrocracking and gasoline blending were commercialized in the 1980s and deployed throughout the HPI and the chemical industry ever since. Control systems engineering has been implemented for mechanical and electrical systems like aircraft and spacecraft since 1960. • Earth’s temperature system cannot be adequately measured or controlled. Mathematical criteria devised in the 1960’s that ensure a system is measurable, observable and controllable are not satisfied. • Mankind has no decision process for properly setting global temperature or CO2 targets, or home thermostats either. The rigorous procedure for optimizing risky tradeoffs for HPI control system setpoints like thermostats was published in HP, December 1996. 28

I FEBRUARY 2010 HYDROCARBON PROCESSING

Ethics.

• Gradual warming is good. Earth’s flora, fauna and humans have flourished since Earth warmed again 18,000 years ago. Humans have experienced 5,000,000 years/50,000 years per cycle = about 100 such cycles. New Yorkers retire to Florida, Canadians to Phoenix, Chicagoans to Hawaii and Germans to Provence. • Taxing energy production is bad. Energy management is basic to human prosperity and well being. Profitable conversion of heat to work since 1780 has created great comfort and wealth for all who know how. Waxman-Markey HR2454 will never work. • India, China, Africa and Russia will continue to produce CO2 from coal, oil and gas, to their credit. Their people will prosper. • Al Gore, at Oxford on July 8, 2009, promoted research to violate the second law of thermodynamics. He condemned power plant and vehicle combustion for wasting 70% of the fuels energy. In 1824 Sadi Carnot proved the maximum theoretical frictionless reversible efficiency is Wo/Qi = 1 – T2/T 1, where Qi is total heat in, Wo is net work out, T 1 is temperature of the heat source (flame, steam) and T2 is temperature of the surroundings (air, cooling water). Great engineers have labored to approach maximum economic efficiency ever since. • Corrupting science is bad. Al Gore promotes spending by governments around the globe to finance his multibillion-dollar venture-capital fund, KPCB, which owns 16 Greentech firms and Google. Providing government grants for fraudulent science research promoting caps on CO2 production is a conflict of interest. I personally found flawed science in peer-reviewed papers in Science and Proceedings of The Royal Society and published my findings in a letter to HP in January 2009. • On April 17, 2009 the US EPA issued instructions for comments on, “Proposed Endangerment and Cause or Contribute Findings for Greenhouse Gases under the Clean Air Act,” as it prepared to declare CO2 a pollutant. It claims current law and court precedent authorize them to do so. Conclusions.

• Knowledgeable environmental engineers support reforestation and efforts to curtail anthropogenic pollutants like SO2, NOx, Bz, CFCs, particulates and surface ozone. They oppose depriving Earth’s flora of their green plant food, choking and starving them for personal gain. I like harmless CO2. I exhale some at 40,000 ppm every 4 seconds. • CO2 and O2 are the basic molecules of the life cycle between Earth’s flora and fauna. The miracle of life photosynthesis reaction should not be tampered with lightly. Starving and choking plants of their food supply would be a monumental crime against humanity, all fauna and flora, the environment and Earth itself. • Since there are no graduate or licensed chemical process control engineers in the UN IPCC, US Congress, Cabinet or Supreme Court, these incompetent engineering groups continue to waste time and money since 1997 attempting the impossible, designing Earth’s thermostat using anthropogenic CO2. No one has controlled the climate of an entire planet. • Climate experts like Richard Lindzen, William Happer, S Fred Singer and ClimateDepot.com are reliable. • Forecast: This article will remain valid beyond 3000 AD. If engineers consider this report good news, ok. I welcome any proof of errors and apologize if I have offended anyone. If we knew what we were doing, it wouldn’t be called research. Albert Einstein. HP

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HPI VIEWPOINT Meeting US energy challenges— Methanol as an alternative transportation fuel John Floren, Methanex’s senior vice president of global marketing and logistics, is an accomplished executive with more than 25 years of diverse international management experience. He manages a global marketing and logistics team of approximately 100 employees and oversees four regional marketing offices as well as waterfront shipping. As a member of Methanex’s executive leadership team and the global leadership council, he provides leadership in the management of the corporation and contributes to the strategic direction of the company. Prior to this role, Mr. Floren was based in Dallas as Methanex’s director of marketing and logistics for North America. He joined Methanex in 2000 in the role of director of information technology and e-commerce before joining the marketing and logistics function. Prior to joining Methanex, Mr. Floren was with Quadra Chemicals Ltd. for 14 years, where he held a number of positions. When he left Quadra to join Methanex, he was the vice president and general manager. Mr. Floren currently sits on two external boards. He has been the chairman of the board of the Methanol Institute since January 2007 and a board member of the Phoenix Academy of Learning Society since January 2009. He is also a member of the Standardbred Canada Audit and Finance Committee. Mr. Floren earned his BA degree from the University of Manitoba in 1987. He is also a graduate of the Harvard Business School, Program for Management Development and attended the International Executive Program at INSEAD in France.

It is widely believed that in coming years, the United States (US) will require increased production of alternative fuels to support an increasing population, a growing economy and greater energy needs. Alcohol fuels like methanol can play a major role in helping the US meet some of these challenges. To be successful in the market, alternative fuels require a large resource base, economic sustainability, consumer acceptance and a positive environmental impact. Methanol contains all these essential elements and is emerging clearly as a transportation fuel that can provide the US with fuel diversity, consumer choice and energy independence. Large resource base. Methanol is a biodegradable liquid fuel at room temperature typically produced from natural gas, although it can be produced from any carbon-based source, including carbon dioxide (CO2), coal and biomass. The opportunity exists for the US to make use of its immense reserves of natural gas, coal and biomass to produce clean-burning methanol that can act as a substitute to petroleum and help consumers at the pumps by offering them more fuel choices and diversity. China already leads the world in the use of methanol as a transportation fuel. This year alone, according to Methanol Market Services Asia (MMSA), China will have blended over one billion gallons of methanol in its gasoline pool. Growth has been rapid in recent years; methanol now represents over 5% of the total Chinese transportation gasoline fuel pool and its use continues to grow. The Chinese government has stated publicly that it considers methanol a strategic fuel and has recently enacted national

methanol blending standards. In many parts of the country, bus and taxi fleets run on methanol blends and retail pumps sell blends of methanol and gasoline. China’s automotive industry is also adapting quickly to this new reality. Several leading Chinese automakers, such as Shanghai Automotive Industry Corp. and Geely Holdings Group, have completed demonstration work on methanol-compatible vehicles for high-methanol blends and have announced that they are ramping up to full-scale production. Economic sustainability. The US can use its own CO2, coal,

natural gas and/or biomass resources to produce methanol and create high-paying jobs for its citizens, reducing its reliance on foreign oil and enhancing its own downstream chemical and energy sectors. Along with methanol’s economic advantages, policy incentives favoring alcohol fuels are also being designed by the US Congress, supporting the view that alcohol fuels are becoming an essential part of the energy mix as a viable alternative fuel. In July 2009, the US House of Representatives passed the landmark American Clean Energy and Security Act, which included provisions to allow the US Secretary of Transportation to require that a certain proportion of new cars built in the US be made “flexible fuel vehicles,” i.e., warranted to operate on alcohol fuels such as methanol and ethanol, as well as biodiesel. Consumer acceptance. Methanol can be blended directly

into gasoline, offering additional fuel choices to US consumers and an alternative to the import of petroleum products. Its unique characteristics include high octane, increased power and increased torque. And, while methanol does have half the energy content of gasoline, it is one of the lowest-cost alternative blendstocks for gasoline. What is more, engine and vehicle tests conducted indicate that vehicles designed to operate on methanol can be 25% to 30% more fuel efficient than regular gasoline-fueled vehicles. Introducing methanol as a gasoline blendstock can easily be done, provided minor modifications are made to the existing fuel infrastructure. In addition, conversion of gasoline vehicles to flexible fuel vehicles capable of running on methanol, ethanol and gasoline would average around $50 to $100 per vehicle. Detractors of methanol often raise its toxicity as a reason not to use this fuel. Methanol, just like other chemicals and gasoline, must be handled according to established guidelines, regulations and codes of practices. Years of methanol use as a transportation fuel, both in North America and in China have demonstrated that, when used appropriately, methanol can indeed be used with a safety profile similar to gasoline and diesel fuel. Positive environmental impact. Methanol’s clean-burning properties and numerous environmental benefits make it a sound environmental alternative. Vehicles with engines that are designed to operate on 100% methanol (made from natural gas) can reduce greenhouse gas emissions by 25% to 30%, compared to regular gasoline-fueled vehicles. Trucks and buses designed to run on 100% methanol emit almost no particulate matter and HYDROCARBON PROCESSING FEBRUARY 2010

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HPI VIEWPOINT notably less nitrogen oxides than their diesel-fueled counterparts. Methanol fuel blending can also reduce urban ozone—a major component of smog. Methanol is also soluble in water and is easier to clean up than gasoline in the event of a spill. It should also be mentioned that methanol has been used for years in the US as an essential feedstock in the production of biodiesel, a clean-burning alternative fuel produced from domestic, renewable resources.

Looking ahead. From this combination of reasons, it is clear why

policy makers are now looking at methanol as a viable transportation fuel that can help the US reduce its reliance on foreign oil and support the need for diverse, clean-burning alternative fuels. Methanol offers many advantages that make it a desirable transportation fuel for US consumers: a wide availability of feedstock, competitive pricing relative to gasoline, ease of introduction into the current fuel infrastructure, and significant environmental benefits. HP

World biodiesel capacity expands Robert R. Starkey holds the position of vice president of fuels for Jim Jordan & Associates, LP, which is an independently owned consultancy and fuels market analyst which provides analysis and intelligence to the global transportation fuels, ethanol and methanol industries worldwide. In this position, he is responsible for the management and direction of the fuels consulting business of JJ&A. This includes responsibility for direction of the fuels team, the weekly fuels publication, client interface on fuels related topics and managing single and multi-client projects. Mr. Starkey has 25 years of extensive experience in sales and trading of petrochemicals and biofuels. His most recent position was president of BPI, an international consulting and trading firm, which included development of ethanol production projects. He holds a BA degree in chemistry from the University of San Diego, completed studies for a masters degree in agricultural biochemistry from the University of Arizona at Tucson, and holds a masters degree in international management from the American Graduate School of International Management (Thunderbird) in Glendale, Arizona. More information on JJ&A and the 2010 World Biodiesel study can be found at www.jordan-associates.com.

Biodiesel use is growing worldwide, driven by many of the same factors faced by all governments: energy security, lessening dependence on imports, improving incomes for farm families and providing jobs. Along with concerns for a clean environment and climate change, these factors have influenced legislation that has been driving the world market for biofuels production and consumption. Early incentives drew investment into the industry that grew at breakneck speed around the world. Global biodiesel capacities have grown from 3 million metric tons to over 36 million metric tons over the past five years. The world consumes about 1.6 billion gallons of biodiesel annually. This is almost a 12-fold increase since 2002; worldwide biodiesel demand is expected to increase five-fold by 2015 to 7.8 billion gallons. Many countries are using Brazil’s sugarcane-toethanol success as a model to turn their own agricultural feedstocks into biodiesel and develop a domestic market, with an eye toward exports. However, the rapid expansion of this market has left this industry overbuilt and ill-equipped to handle many of the present challenges it now faces. US market. Biodiesel producers in the US have been operating well below their production capacity since their primary export outlet was cut off in March 2009. Heavy import duties placed by the European Union (EU) on US biodiesel stopped exports, which accounted for 75% of US production. Prior to the import duties, US production capacity was not operating even close to 32

I FEBRUARY 2010 HYDROCARBON PROCESSING

the installed capacity. The National Biodiesel Board places current capacity at 2.68 billion gallons. According to the Energy Information Administration data, no more than 680 million gallons have been produced in a single year. The Renewable Fuels Standard (RFS-2) called for 500 million gallons to be blended into the diesel fuel supply in 2009. However, with the lack of blending rules and assignment of obligations, obligated parties and fuel blenders are uncertain how to comply with this mandate. The EPA, which is responsible for creating the rules to implement the mandate, is allowing the 2009 blending requirement be rolled into 2010 since it has not yet issued the rule. At the writing of this article, there are still no rules for blending; 2010 sees the addition of 600 million gallons to the mandate. The EPA rules are under review by the Office of Management and Budget. To add insult to injury, the $1.01 per gallon tax credit for blending biodiesel expired on December 31, 2009. Congress failed to act on this tax credit extension before going on break for the holidays. Although this creates considerable concern for biodiesel producers and investors, Senate Finance Committee Chairman Max Baucus (D-Mont.) and Chuck Grassley (R-Iowa) plan to move on an energy tax extenders package early in 2010 which will likely include retroactive extension of the biodiesel tax credit. However, many biodiesel fuel producers will have to shut down until the tax credit is reinstated. This could create a shortage of biodiesel for some existing private and government fleet customers who have made a policy commitment to use alternative fuels. European market. As precarious as the biodiesel situation appears in the US, it is not that different from biodiesel’s situation in the rest of the world. Because of a poor compliance record, the EU is in the process of revamping its earlier directive of 5.75% biofuels in the transportation sector by making this biofuels target mandatory for 2010. This is a positive move for the industry in Europe, but the EU is relying on each member country to develop an individual national plan for meeting this goal. Performance by Germany, France, Poland and Spain have been in line with the original directive, but the remaining member countries have a long way to go, and they will most likely not reach the goal until much later in the decade. As in the US, Europe’s biodiesel production capacity exceeds demand; yet, cheaper imports from Argentina, and those formerly from the US are allowed to displace domestic production. Germany’s government supported the biodiesel industry by eliminating excise taxes at the retail pump. However, starting in 2008, the tax incentive was removed, placing biodiesel at the same price as petroleum diesel, and consequently, demand for biodiesel dropped. This shut down domestic production and left the biodiesel distributors with the only option of importing

HPI VIEWPOINT cheaper biodiesel to allow biodiesel blended fuels to compete with petroleum diesel. Asia-Pacific market. In Asia-Pacific, the two largest palm oil producing and exporting countries, Indonesia and Malaysia, aggressively embraced biodiesel production and authorized rapid development. In 2007, commodity prices for palm oil began a rapid climb, making biodiesel too expensive to produce for the local market, which, in turn, idled much of the capacity. When crude oil prices collapsed in 2008, petroleum diesel became much cheaper than palm-oil-based biodiesel. Palm oil exports continue, but domestic demand for biodiesel has not recovered. Europe and the US have prejudiced the use of palm oil because of concerns over land use changes and rainforest destruction. This may, in fact, reduce price pressures for this feedstock in Asia-Pacific and allow for biodiesel exports to China, Japan and South Korea. India is one of the largest consumers of Malaysian palm oil, and this nation built biodiesel capacity with the intention of using cheap palm oil imports. However, in 2008, India banned

the usage of edible oils for biodiesel production due to high costs and concerns over edible oil shortages for food use. Instead, many biodiesel companies are switching to other sources, including jatropha oil. By some estimates, over a million hectares have been planted in India, with commercial oil production expected in 2013 to 2015. Global view. Worldwide, the biodiesel industry is struggling

with high feedstock costs and overcapacity. If not for numerous government mandates around the world, this industry would have most probably disappeared after the commodity price increases in 2007 and 2008. The big question is whether government support and mandates will last until competitive feedstock alternatives, like jatropha and algae, can enter the market. For additional information on the global biodiesel picture, Jim Jordan & Associates (JJ&A) has prepared a global biodiesel study that examines production capacities, supply and demand balances. We include a forecast for production and consumption from 2010 through 2014. HP

Ethanol: Growth of a clean and renewable energy source José Carlos Grubisich is president of ETH Bioenergia. He graduated in 1979 with a degree in chemical engineering from Escola Superior de Química Osvaldo Cruz, and he holds an MBA in Advanced Management issued by INSEAD, France. Mr. Grubisich began his career at the RhônePoulenc Group, where he held several positions in Brazil and abroad. His main roles were president of Rhodia Brazil and Latin America from 1997 and vice-president of Rhodia Worldwide from 2000. Mr. Grubisich became a member of the group’s Executive Committee. In January 2002, he headed the integration process of the companies that would become Braskem, formed officially in August of the same year. As the company’s president, he headed the process that would consolidate Braskem as a world-class Brazilian petrochemical company. For six consecutive years (2002–2007), he was awarded the “Prêmio Executivo de Valor,” by the Jornal Valor Econômico newspaper. In July 2008, he took over as president of ETH Bioenergia.

he global energy matrix is changing. Leaders and scientists in the developed nations are focusing on the impact of carbon dioxide (CO2) emissions and the effects of the climate change caused by man. The civilization based on crude oil extraction is making way for renewable energy sources and this is reversing a trend that has dominated most of the 20th century. Such a change occurs at the same pace in which the negative effects of global warming are felt, and global powers such as India, China, the US and the European Union are taking action to introduce new clean energy sources into their domestic energy matrixes.

Brazilian market. Brazilian ethanol production doubled to 24 billion liters between 2006 and 2008. This growth trend is forecast to continue through 2015 and to pass the 50 billion-liter mark for ethanol production. The transformation of the Brazilian ethanol market occurred with the introduction of flex-fuel vehicles, which gave consumers the chance to decide when to use

ethanol or gasoline. Increasing income levels and growing credit lines for automobile purchases enabled flex-fuel-powered vehicles to dominate the market. Note: Nearly 90% of the cars currently sold in Brazil are “flex,” which will further increase domestic ethanol consumption. The Brazilian tradition in the sugar and ethanol sector gave this country an enormous advantage with regard to practices and technologies. As the global demand for renewable fuel resources increases, sugar-cane ethanol appears to be one of the most feasible options, and this is an excellent commercial opportunity for Brazil. In addition, it is important to stress that the sugar and ethanol sector is at a new level, supported by a pyramid support of professionalism, technology and sustainability. Renewable resources. Brazil must make the most of this opportunity to consolidate its position as the global leader in this sector, as 46% of its domestic energy comes from renewable sources. This nation is in the best position to become a world leader in the production and export of agricultural products, including commodities and biofuels. When considering the positive estimates on ethanol consumption over the next few years and the five billion liters of ethanol exported in 2008, the moment for the Brazilian production to enter the US market, with President Obama’s sustainability agenda, and the European market, is becoming closer. Such a scenario is opening the market to investment from large corporations such as ETH Bioenergia. In less than two years, ETH developed a business model that combines competitiveness and sustainability, based in three production areas: the states of São Paulo, Mato Grosso do Sul and Goiás. This model allows synergy among all units, production scale and competitive costs, through the use of cutting-edge technology in the production process and professional qualification of its labor. Although it is a new player in the market, ETH Bioenergia will double its production crop in the 2009/2010—it is growing at an accelerated rate. This is an ambitious initiative that moves forward at a firm pace and responds to the challenges brought by the need for clean and sustainable energy at a global scale. HP HYDROCARBON PROCESSING FEBRUARY 2010

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CLEAN FUELS

SPECIALREPORT

Consider new methods for bottom of the barrel processingâ&#x20AC;&#x201D;Part 1 Advanced methods use molecule management to upgrade heavy ends M. MOTAGHI, KBR, Houston, Texas; and K. SHREE and S. KRISHNAMURTHY, KBR, New Delhi, India

arge price differentials between light, sweet crudes vs. heavy, sour crudes have created strong incentives for refiners to lower costs by incorporating as much heavy crude blends into the refinery processing in scheme as it can tolerate. Several refineries have little or no bottoms processing capabilities, thus yielding large volumes of high-sulfur fuel oil (FO). New bunker-fuel legislation and pending carbon-footprint initiatives create the need to further upgrade refinery resid products, both for expansion and enhancement reasons. Several case histories will cover the staggered investment options to produce premium road asphalt and solid fuels (Fig. 1). Changing crude diet. Recent economic and geopolitical global uncertainties are factors in the steep rise in crude prices, thus affecting refinery operations worldwide. Refiners are finding themselves faced with the difficult quest to search for crude blends to maximize margins. While refinery margins will continue to be dictated by processing heavier, more sour crudes, the dramatic increase in residuum content from 10% in light sweet crudes to 50% in extra heavy crudes poses interesting challenges, while presenting some unique opportunities. The bulk of the global operating refineries have little or no residuum processing capabilities and produce large volumes of high-sulfur FO (HSFO) and bunker fuel. As demand shifts to natural gas, FO demand is expected to drop adversely, thus affecting FO prices in the future (Fig. 2). This situation is only expected to worsen as refiners face regulatory pressures ranging from new maritime bunker fuel specifications

to carbon dioxide cap-and-trade and carbon footprint limitations. As the world moves towards cleaner bunker fuels, finding alternative means to upgrade bottom of the barrel streams will become increasingly important. In parallel, the demand for jet fuel and diesel is expected to grow, and many refiners are actively focused on shifting demand from motor gasoline to diesel, while still analyzing all available options. The price differential between diesel and gasoline is likely to be sustained over the long-term and is validated by the billions of dollars of investment announcements by the major international and national oil companies toward dieselization. In markets dominated by fluid catalytic cracking (FCC)-based refineries, this need to increase distillate production has taken on a new dimension that will impact long-term refining margins. Changes in C/H ratios. In its simplest form, refining is the process of changing the carbon-to-hydrogen (C/H) ratio of naturally occurring crude oils. Thus, at the molecular level, the operation of all 650 refineries in the world is essentially targeted at converting high C/H ratio feedstocks into high hydrogen to carbon ratio for transportation fuels. This ratio change between the crudes and products can only be accomplished through two broad processing routes: carbon rejection and hydrogen addition. While profitability can only be sustained by economically converting the large volumes of residuum into high-value transport fuels, these objectives must be accomplished in a difficult business climate that mandates applying low-investment solutions while optimizing usage of existing refinery resources.

FO crack margin, $/bbl

-10

-20 2000

2004

2008

Source: Singapore margins w.r.t. Dubai crude from Platts

FIG. 1

The residuum oil supercritical extraction unit at the Navaho Refinery.

FIG. 2

FO cracking margins 2000 to 2008.

HYDROCARBON PROCESSING FEBRUARY 2010

I 35

SPECIALREPORT

CLEAN FUELS

Carbon rejection. Carbon rejection is favored by low crude

are limited in processing lighter, low-metals, low-sulfur residues. Visbreakers are essentially a means of improving viscosity of the residuum so as to minimize the addition of valuable distillate boiling range cutter stock to meet FO specifications. In addition, with the increasing shift in product demand from motor gasoline to diesel and jet fuels, these technologies will no longer be the first choice. Residues from heavy crude oils contain high concentrations of sulfur, complex hydrocarbons and heavy metals such as nickel and vanadium. Due to the nature of these residues, delayed coking technology is the most commonly used carbon-rejection technology. However, this process produces highly hydrogen-deficient, unstable products that require further processing and yields coke residue—a high C/H ratio molecule, as a byproduct. Although, coking is a mature technology with low implementation risks, in a carbon footprint future, this technology is likely to Pfl j\\ X ]XZ`c`kp% face stiff environmental and regulatory resistance even in the face of lower crude prices. N\ j\\ Zfdgc\k\ N\ j\\ Zfdgc\k\ jX]\kp% High-sulfur petroleum coke prices are distressed and as is evident in the Canadian inland environment, coke is just being piled up in large quantities with no real economic outlet. This trend cannot be sustained in the long run.

prices and high hydrogen prices, when it is economical to reject the residuum as petroleum coke, while producing the required transport fuel volumes by incremental crude oil processing. During these processes, the bulk of feed contaminants are rejected with the carbon into the coke, rather than into the liquid products. Traditional carbon rejection techniques include FCC, resid fluidized catalytic cracking (RFCC), delayed coking (DC) and visbreaking (VB). FCC and RFCC in combination are widely used carbon rejection technologies to convert the high-boiling, high-molecular weight hydrocarbon fractions into more valuable gasoline, olefinic gases and other products. However, due to the nature of the process, they

Hydrogen addition. Conversely,

?lek\i J`k\ J\im`Z\j f]]\ij X mXjk befnc\[^\ Xe[ k_\ Y\jk k\Z_e`hl\j ]fi k_\ Ô\c[ Xjj\dYcp f] YcXjk$i\j`jkXek le`kj% =ifd Yl`c[è^ j\kkè^ kf fg\iXkè^ ]XZ`c`kp èk\îXk`fe# ?lek\i J`k\ J\im`Z\j gifm`[\ k_\ ]lcc jg\Zkild f] Zfdgfe\ekj ]fi Xe \]ÔZ`\ek# Zfjk$\]]\Zk`m\ fe$j`k\ j\klg% ?lek\i J`k\ J\im`Z\j _Xj Xcc f] k_\ ZXgXY`c`k`\j ]fi jlZZ\jj]lc j`k\ Xjj\dYcp# èZcl[è^1 I\j`[\ek k\Z_e`ZXc befnc\[^\ G\idXe\ek Xe[ c\Xj\[ Yl`c[è^ fgk`fej KfkXc gifa\Zk \o\Zlk`fe D\Z_Xe`ZXc èk\î`kp èjg\Zk`fej Iflkè\ Xe[ gi\m\ekXk`m\ dXèk\eXeZ\ <ok\e[\[ Zljkfd\i nXiiXekp <]]\Zk`m\ lj\ f] Zljkfd\iËj èk\ieXc i\jfliZ\j

J\\ jX]\kp ]ifd fli j`[\ Ç ?lek\i J`k\ J\im`Z\j b\\gj pfl gifk\Zk\[ ]ifd k_\ flkj`[\ `e%

J@K< J<IM@:<J

_lek\iYl`c[`e^j%Zfd 36

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)/(%+,)%0/''

hydrogen addition is favored by high crude prices and low hydrogen prices when it is more economical to upgrade the residuum to transport fuels, while maximizing the transport fuels production from the base crude capacity. Thus far, residue-hydrogen addition technologies have focused on fixed-bed hydrocracking as against ebullated-bed hydrocracking and slurry-phase hydrocracking, the former requiring periodic shutdowns to regenerate catalyst. Ebullated-bed processes are continuous and produce higher levels of liquid fuels (no coke). But they are unable to achieve complete resid conversion and still produce 20%–30% of heavy-resid product. Ebullated beds have also been prone to high operating costs, and have sometimes been plagued with poor operability. The quality of liquid products, although improved over coking, still requires secondary processing to produce clean fuels. The inability to achieve near complete conversion requires further processing of unconverted resid. As a result, ebullated-bed technologies have not achieved huge deployment, which, when coupled with the high capital cost, makes them the least robust at low-oil price scenarios. Slurry-phase hydrocracking. The recent flurry of activities indicates the advent of slurry-phase hydrocracking into the market place. This technology adopts high operating pressures and can achieve near complete conversion of the residuum while producing finished saleable products. While the pro-

CLEAN FUELS jected economic conditions of high crude prices and low gas prices are ideal for investment in these high-pressure/high conversion technologies, the capital investment requirements may have some dampening effect for immediate widespread adaptation. Principles of molecule management. In the background of high volatility within the markets, which is expected to continue, the optimal solution may require most refiners to adopt a combination of carbon rejection and hydrogen addition processes. This is especially true in an environment of uncertain refining margins where the size of capital investment can come under a higher scrutiny over traditional project â&#x20AC;&#x153;return on investmentâ&#x20AC;? criteria. It is in this context that a staggered investment option involving the ability to achieve partial benefits at lower initial investments, while preserving options for incremental benefits with higher investment in the future, gains increasing importance. Refiners are often constrained by the need to convert a defined crude slate to a set product slate without realizing the change required at the molecule level and the cost associated with such conversion. Refiners can also be blinded by the compelling need to produce traditional refinery products such as transport fuels and petrochemical feedstock from every barrel of crude, often expending substantial capital in the process, while ignoring economic synergistic opportunities that may exist with other nontraditional industrial applications. The principles of molecule management dictate that the best economics are derived by capturing the highest value of every molecule present in naturally occurring crude oils at every point in the process. When viewed in this context, it is evident that it will be prudent to analyze the residuum fraction not only by the traditional barometers of boiling range and gravity, but by molecular speciation. While distillation based separation schemes for the virgin crude fractions are economical and adopted almost universally, in almost all cases, the volume and quality of the residuum is essentially determined by the quality of the vacuum gasoil (VGO) fraction and the ability to process this fraction through conventional hydroprocessing or catalytic cracking conversion units. In most cases, the limiting factor is the metals content or the Conradson Carbon Residue (CCR) in the GO. The residuum volume and quality is, by balance, a reject defined by GO quality, and is characterized as black oil. By conventional wisdom, this stream is either removed as FO or asphalt, or is subject to thermal conversion processes for upgrading. While it is a well-established fact that hydro or catalytic conversion of the heavy gasoil (HGO) fractions will result in substantially better yields and qualities of transport fuels (gasoline, jet fuel and diesel) than ther-

SPECIALREPORT

mal conversion processes, and the incentive to maximize this fraction of the crude exists, operating economics are substantially influenced by the incremental concentration of impurities (metals and CCR) in the feed due to their impact on conversion unit catalysts. When analyzing the residuum at a molecular level, it will be evident that a substantial volume of higher boiling range white oil molecules worthy of effective catalytic upgrading are present in this fraction, which by conventional methods, are rejected as black oil products or subjected to thermal conversion processes. This phenomenon is essentially caused by limitations in distillationbased separation processes where the lowest boiling point species of the undesirable impurity is the determining factor in the vol-

Select 155 at www.HydrocarbonProcessing.com/RS 37

SPECIALREPORT

CLEAN FUELS

umes of the GO and residuum derived. This impacts the refinery product slate and economics negatively. It is, therefore, essential to look at supplemental alternate separation technologies to effectively extract these higher boiling white oil molecules from the residuum. The options become obvious when analyzing these molecular species, and it is clear that these undesirable impurities are essentially asphaltenic or resinic in nature and can be separated by solubility driven processes.

from the resinic and asphaltenic molecules contained in the residuum. SDA uses a paraffinic solvent, which by molecular structure (like dissolves like) preferentially dissolves paraffinic and naphthenic molecules while rejecting the aromatic-rich molecules in the pitch (Fig. 3). Although light-paraffin solvent-based deasphalting is often referred to as a metals or CCR rejection process, in essence, it is an aromaticsrejection technology. The reject contains complex aromatic molecules that are the least soluble in paraffinic solvents, are highly hydrogen deficient and contain a majority of polars (metals and CCR) that are White oil molecule management. The solution involves least desirable when fed to a hydro or catalytic conversion unit. applying the solubility-based physical separation process, solvent Conversely, the extract or the deasphalted oil (DAO) contains deasphalting (SDA), wherein the saturates can be effectively separated essentially the saturates, with very low metals and CCR, making this a white oil ideal for conversion processes. Also, the top barrel of DAO derived from the residue via solvent extraction by all measures, will always represent a superior conversion feedstock to any secondary unit when compared to the last barrel of VGO derived from a distillation process. When examining the suitability of feedstocks to a conversion process, the inherent molecular content (as opposed to boiling range) and their impact on conversion unit performance will become obvious. For example, a close examination of the Watson K factors of feeds derived from GO fractions ® Potassium vs. Zyme-Flow and resids from staple Arabian crudes to an FCC unit will reveal that the DAO derived Permanganate UN657 from a solvent extraction process brings in resid boiling range white oil molecules that are not otherwise achievable by conventional distillation based processes (Table 1). The DAO, rich in paraffins and saturates, is an excellent feed to the FCC. Also, due to the inherent nature of the aromatic rejection associated with solvent-based separation processes, the DAO will always have a higher Watson K than the corresponding VGO from the same crude. As listed in Table 2, the FCC yields and Powerful Safe product qualities derived from a DAO based Zyme-Flow® UN657 eliminates Zyme-Flow® UN657 is nontoxic, H2S and removes pyrophoric biodegradable, compatible with Easier to process downstream iron sulfides, benzene and all metallurgies and friendly to hydrocarbons. It degasses, waste water treatment plants. de-oils and reduces waste in What’s more, it’s an all-aqueous, a single step – typically cutting VOC-free solution that creates decontamination time by 50% no hazardous waste for disposal. or more. Feed DAO Zyme-Flow ® UN657. Part of a full range of innovative decon chemistries and planning services from ULI.

Asphaltene

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Asphaltenes w/CCR, metals, nitrogen, etc. FIG. 3

Solvent de-asphalting process and product streams.

CLEAN FUELS TABLE 1. Watson K of FCC feed fractions Feedstock source

Atmospheric resid

VGO

Propane DAO

Butane DAO

Coker GO

Arabian light

11.60

11.68

11.81

11.74

11.4

Arabian heavy

11.44

11.62

11.86

11.78

11.4

feed (Mid Continent) is very similar to the VGO yields from the same crude in spite of processing heavier molecules derived from the resid boiling range. The alternate option would be to reject these valuable resid white oil molecules to lower value black oil products, or to process through a lower yield thermal conversion process while reprocessing the derived coker GOs in the FCC. The understanding of the value of these high hydrogen content resid molecules and their disposition is a critical data point in determining refinery yields and economics. An understanding of the metals and CCR content in the DAO stream and their impact on conversion unit catalysis is critical to determining the extractable volume of white oil molecules contained in the residuum. While metals are inherently addressed by demetalization catalysis in a hydrocracker, or by FCC catalyst consumption, and the economics are easy to calculate, the understanding of the CCR is more complicated (Table 2). Notice that the CCR content of DAO is nearly 10 times higher than the VGO, yet the yields and product qualities and coke make are still comparable. The obvious inference is that the CCR in this case does not impact the conversion simply by virtue of the molecules that contribute to it. In its simplest form, CCR is the residue derived from a test wherein a heavy hydrocarbon is subject to a temperaturetime exposure. While the temptation to use CCR as a parameter to establish conversion unit performance exists, a molecular level examination will highlight significant dif-

SPECIALREPORT

ferences. In short, all CCRs are not the same. While the heptane insolubles have an almost direct permanent deactivation effect on downstream process catalysts, the other CCR molecules have a far less telling effect. The CCR derived from a distillation-based separation process will contain substantial C7 insolubles, while that in a DAO derived from a solvent extraction, by definition, should be nondetectable. This distinction can be clearly demonstrated by operating performance differences between units processing atmospheric or vacuum resids and those processing DAOs with identical CCR content. This difference is often not well understood by refiners while correlating the impact of CCR on their unit performance.

Low tolerance for pump problems? No problem.

TABLE 2. FCC yields from Mid-Continent VGO/DAO and CGO

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100% DAO

100% VGO

100% CGO

API

19.2

24.7

19.0

S, wt%

0.79

0.75

CCR, wt%

3.9

0.39

Less than 1

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Ni + V, PPM

Less than 1

Proven benefits include:

FCC yields, wt% Conversion

80.3

81.05

63.2

C2–

4.86

3.65

1.49

Total C3’s

6.37

6.80

4.60

Total C4’s

10.30

11.76

8.87

Total gasoline

48.98

52.12

40.16

Total cycle oil

19.70

18.95

35.78

Coke

9.79

6.72

6.00

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TABLE 3. Commercial data for conversion Supercritical SDA

After conversion to residuum oil supercritical extraction

Feed rate, bpsd

7,000

10,000 (up to 15,000)

Solvent ratio vol/vol

4.5–6

5–6

Energy MMBtu/bbl

65–75

70–85

Deasphalted oil quality Yield, LV% Asphaltene, (C7 insols) ppmw CCR, wt%

200–800

< 25

12–13

9–11

With trays

With advanced intervals

With mesh pad

With advanced intervals

Table 3 lists commercial data published by a US refiner that shows the before and after conversion of a non-supercritical SDA unit to an advanced residuum oil supercritical extraction technology with advanced internals. As evident from the data, the addition of new internals and adoption of residuum oil supercritical extraction methods achieved a higher throughput, higher yield and a better DAO product quality. All this was achieved at similar operating conditions with significantly lower specific energy consumption. Note, the decrease in asphaltene content of the DAO to nearly below detectable limits, conforming the assertion that the presence of C7 insolubles in the DAO is an artifact of entrainment that can be controlled by good technology and internal design features. HP

High pressure vessels Extractor DAO separator

While the DAO must theoretically contain no C7 insolubles, an examination of the DAO derived from conventional SDA processes would often indicate levels ranging from 300–1,000 ppmw. This can only be explained by the phenomenon of entrainment. With the advent of state of the art structured-packing based internals into the solvent deasphalter separator vessels in 1995, the level of C7 insolubles is now controlled at or below 100 ppmw in most deasphalting units and below 25 ppmw in several high performing units. As an added benefit, the process unit can now operate at about twice the phase rates of conventional separators and provides about twice the mass transfer efficiency of conventional extraction contacting devices.

I FEBRUARY 2010 HYDROCARBON PROCESSING

Next month. In Part 2, the authors discuss new supercritical extraction methods that can be applied to optimize molecule management of residuum. Mitra Motaghi is an associate with the KBR refining technology business unit in Houston Texas, with specific focus on resid and hydroprocessing technologies. She holds an MS degree in chemical engineering from Texas A&M, Kingsville, Texas.

Kanu Shree is an associate with the KBR refining technology business unit in New Delhi, India, with specific focus on resid and hydroprocessing technologies. She holds a BS degree in chemical engineering from the Indian Institute of Technology, New Delhi, India. Sujatha Krishnamurthy is an associate with the KBR refining technology business unit in New Delhi, India, with specific focus on resid and hydroprocessing technologies. She holds a BS degree in chemical engineering from Anna University, Chennai, India.

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CLEAN FUELS

SPECIALREPORT

Designing vacuum distillation for deep-cut bitumen service To process bitumen-derived residue, success depends on the engineering design concepts and methods applied M. GRANDE, Fluor Canada Ltd., Calgary, Alberta, Canada

nergy security issues have refiners seeking heavier crude oils and bitumen (oil sands) as feeds for their operations. Several issues require resolution for a successful deep-cut design or revamp of a vacuum distillation unit (VDU) to process bitumenderived atmospheric residue (AR). Success depends on the design concepts and methodologies applied. A wet VDU design may not only be optimal but also necessary depending on the heavy vacuum gasoil (HVGO) yield cutpoints. This case history outlines design concepts and methodology that can be applied to a VDU to successfully producing desired gasoil (GO) yields for reliable conversion. Deep-cut separation. Upgrader/refinery configuration evalu-

ations may indicate benefits in recovering more of the heavy distillates as HVGO from bitumen-derived AR. The optimal cutpoint between HVGO and vacuum residue (VR) depends on the feed contaminates distribution, refinery/upgrader configuration and product requirements. The optimized cutpoint should consider the downstream HVGO conversion unit that dictates the degree of fractionation required vs. which is realistically achievable. This requires a better understanding of the distribution of contaminates/ properties for the distillate boiling range surrounding the selected cutpoint. For evaluations that indicate benefits in recovering more of the heavy distillates, a deep-cut performance is required for the associated new or revamped VDU that usually entails a sharper fractionation between the HVGO and VR products. To provide reliable deep-cut performance, a sound process design is required that evaluates the three main process parameters associated in recovering distillate material: temperature, pressure and steam. This requires a comparison of a wet VDU design (with stripping and velocity steam) vs. a damp VDU design (with only velocity steam). Particularly for VDUs in bitumen service, this comparison should include other key process design considerations based on the selected HVGO yield cutpoint and fractionation required. These considerations typically are the permissible VDU fired-heater coil outlet temperature (COT ) imposed due to the lower thermal stability of the feed, VDU column feed superheat condition, flash zone de-entrainment effectiveness, and wash section de-entrainment effectiveness as well as fractionation efficiency, stripping section fractionation efficiency and the resulting VDU column diameter. Case study. The feed basis for this evaluation is bitumen-

derived AR with a yield cutpoint of 343째C true boiling point

(TBP). The objective is to achieve an HVGO yield cutpoint of 524째C TBP resulting from a relatively sharp fractionation between HVGO and VR products. An optimum VDU column diameter esstimated for achieving this objective with the provided example is labeled as DOptimum . This is the diameter required for the portion of the column that contains the HVGO and wash sections as well as the flash zone. The DOptimum is the reference basis used for illustrating required VDU column diameters applied in the figures throughout this article. The maximum recommended VDU fired heater COT is labeled as COTMax . This is the reference basis applied for illustrating predicted COTs in the figures throughout this article. Higher COTs are illustrated for information purposes only. KEY PROCESS DESIGN PARAMETERS

Key process design parameters include: VDU fired heater COT. Due to thermal stability concerns of bitumen derived AR, limiting the peak film temperature and minimizing oil residence time in the VDU fired heater is imperative. The peak-film temperature should consider the bitumen source and production method. The typical production methods are either non-mined bitumen that is produced with steam assisted gravity drainage and mined bitumen that is either paraffinic-froth treated or naphthenic-froth treated. With the permissible peak-film temperature established, the VDU fired heater COT may be maximized with suitable process design considerations and investment in the fired heater. This sets the temperature limit for distillate material recovery and is labeled as COTMax . To maximize run length of the VDU fired heater, it should be designed to suppress coke formation by minimizing the residence time and film temperature that the AR is exposed to. This includes a double fired-heater design with velocity steam and allowing for a higher pressure drop. The double fired-heater design provides a higher average, but lower maximum heat flux, which reduces the film temperature and residence time of the AR. The velocity steam, with the higher pressure drop allowance, provides a higher average but lower maximum mass flux, which further reduces the residence time of the AR. The velocity steam also maximizes vaporization of AR for a given COT. All of these considerations will minimize the coking potential and, consequently, the amount of cracked-gas production that loads the VDU overhead ejector system. HYDROCARBON PROCESSING FEBRUARY 2010

I 41

SPECIALREPORT

CLEAN FUELS

Consideration should also be given to recycle spent wash oil drawn from the bottom of the wash bed to the VDU fired heater. This minimizes COT (but increases heater duty) because the recycled spent wash oil minimizes the amount of heavier material that is vaporized to supply the wash oil for a desired HVGO yield. The basis for the examples provided is recycled spent wash oil.

to achieve two stages of fractionation. This will improve fractionation efficiency to maintain a reasonable packed-bed height and de-entrainment ability while maintaining fouling resistance in the lower portion of the bed. The basis for the examples provided is two stages of fractionation in the wash section with a maintained overflash of 0.2 gpm/ft2 (0.5 m3/m2 hr).

VDU column wash section. The presence of metals in the HVGO product results from vaporizing heavy organo-metallic compounds and entrainment of VR into the HVGO draw. The metals increase fluid catalytic cracking (FCC) catalyst makeup rate and poison hydroprocessing catalyst. Accordingly, metals concentrations must be controlled within limits. The entrainment effect can be minimized by proper sizing of the column flash and wash sections, transfer line and inlet device. The diameter for the packed sections of a VDU column is typically sized with a CS factor of 0.38 ft/s (0.12 m/s) and is the basis for the examples provided. The vaporization effect is only controlled by limiting the HVGO yield to limit volatilization of organo-metallics. However, since distribution of the organometallics seems to be highly nonlinear, rectification in the wash section of the flash zone and (for the wet VDU design) stripping section vapors minimizes the heavy tail of the HVGO product and, therefore, minimizes the amount of metals for a given HVGO yield. This may also result in a reduction of nitrogen and carbon residue content, which may also be of benefit for downstream hydroprocessing units. To minimize the listed effects, the wash section must not only de-entrain VR and (particularly for AR derived from mined bitumen) ultra-fine solids, but also rectify the flash zone and stripping section vapors. This is accomplished by providing sufficient wash oil (reflux) to a packed bed that has fouling resistance and fractionation ability. The wash oil supply is typically set to maintain a minimum overflash of 0.2 gpm/ft2 (0.5 m3/m2hr). Overflash is defined as the true reflux that exits the wash section, excluding entrained VR liquid, as opposed to the actual spent wash oil that contains entrained VR liquid. A minimum overflash is required to achieve proper wetting of the wash bed that will avoid potential dry spots and associated coking while maximizing HVGO yield. For the packed bed, consideration should be given to lower layers of grid followed by upper layers of a higher specific surface area structured packing

Feed superheat condition. Due to the two-phase flow

COT (COTMax +/- °C)

COT @ 20 mm Hg(A) w/1 wt% steam

-5

Duty @ 36 mm Hg(A) w/2 wt% steam

-10

425 400 375

-15

350

-20

Duty @ 20 mm Hg(A) w/1 wt% steam

325

-25

300

-30

275

-35 0

FIG. 1

450

10 15 20 25 Superheat ΔP,% of total ΔP

VDU fired heater sensitivity to feed superheat.

I FEBRUARY 2010 HYDROCARBON PROCESSING

250 35

120

250

110

Diameter @ 20 mm Hg(A) w/1 wt% steam

100

230 210

Diameter @ 36 mm Hg(A) w/2 wt% steam

190

Wash oil @ 20 mm Hg(A) w/1 wt% steam

80 70

Wash oil @ 36 mm Hg(A) w/2 wt% steam

60 50

170 150 130 110

40 0

FIG. 2

10 15 20 25 Superheat ΔP, % of total ΔP

VDU design sensitivity to feed superheat.

90 35

Wash oil rate, bbl/thousand bbl feed

475

Diameter, % DOptimum

COT @ 36 mm Hg(A) w/2 wt% steam

Heater duty, kW/thousand bpsd feed

regime and high velocities in the VDU column transfer line, literature suggests that the column feed material does not maintain thermodynamic equilibrium as the pressure decreases along the length of the main transfer line.1 One theory is that the pipe fittings/routing local to the exit of the VDU fired heater may assist in maintaining equilibrium by providing sufficient mixing. However, along the main straight section of the transfer line where the diameter is largest, the phases may become a nonequilibrium mixture. This results in a superheated vapor phase that enters the wash zone and vaporizes more of the wash oil. Consequently, to ensure that the overflash is achieved to prevent wash-bed dry out and subsequent coking, a higher wash oil rate to the wash section may be required than predicted by an equilibrium-based model. To maintain yield and produce the increased supply of required wash oil, the required distillate lift is increased by adjusting one or more of the three main process parameters mentioned earlier. Figs. 1 and 2 illustrate the effect of superheated vapor feed conditions on a wet VDU design over a range from full equilibrium (no superheat pressure drop) to approximately the last third of the total transfer line pressure drop contributing to superheat. These figures are based on maintaining an HVGO yield cutpoint of 524°C TBP, a flash zone pressure (FZP ) of 20 mm Hg(A) with 1 wt% stripping steam and 36 mm Hg(A) with 2 wt% stripping steam based on the VR product. These results indicate that wash-bed dry out and subsequent coking, or, alternatively, an undesirable reduction in the HVGO yield, are possible if the design does not consider the effects of a superheated vapor feed condition. Figs. 1 and 2 also indicate that, as pressure and stripping steam rates are increased, the sensitivity to the superheat condition of the vapor feed diminishes. The basis for the examples presented in this article is that a small percentage of the total transfer line pressure drop is associated with the main transfer line header that generates superheated vapor.

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CLEAN FUELS

For both the wet and damp VDU design bases, velocity steam is introduced into the tubes of the radiant section of the VDU fired heater. For the examples presented, a velocity steam rate of 0.2 wt% of fresh charge to the fired heater is applied for both design types. The wet design also includes stripping steam injected into a stripping section of the VDU column. Stripping steam considerations. Stripping steam per-

forms the function of stripping the flash zone liquid of HVGO, which allows less distillate material vaporized from the VDU fired heater for a desired HVGO yield. The amount of stripping steam required for a desired stripped HVGO lift is minimized by maximizing the distillate lift to stripping steam ratio. This is achieved by maximizing the fractionation efficiency of the stripping section. The basis of the examples provided is two stages of fractionation for the stripping section. The shift of distillate lift from the fired heater to the stripping section allows for a lower fired heater COT and/or a higher FZP, whichever is determined as most beneficial. Adjusting these process parameters requires consideration of the thermal stability of the bitumen-derived AR and the desired HVGO yield. These adjustments have significantly different impacts on the operating and capital expenditures for the VDU. If the FZP is maintained to minimize the COT, a relatively larger diameter is required for all sections of the VDU column due to the presence of stripping steam increasing the vapor volumetric flow profile. Sour-water production from the VDU will also increase for this case due to the addition of stripping steam. Fig. 3 illustrates the above relationship (the diameter for the portion of the column containing the HVGO and wash sections as well as the flash zone is used) to achieve an HVGO yield cutpoint of 524°C TBP. A FZP of 36 mm Hg(A) is maintained for this example with a large portion of the illustrated range residing above the recommended maximum COT. Conversely, if the COT is maintained to maximize FZP, it will be demonstrated that, even with the stripping steam present, a smaller diameter is required for all sections of the VDU column due to a lower vapor volumetric flow profile. This effect of a diameter reduction is limited to a determinable minimum diameter. Sour-water production may also decrease for this case, even with the stripping steam present, due to the reduced motive steam required for the optimized FZP.

105

102 99

93 90 0.0

FIG. 3

COT

Diameter

0.5

1.0 1.5 2.0 2.5 Stripping steam, wt% on VR

VDU design sensitivity to stripping steam.

I FEBRUARY 2010 HYDROCARBON PROCESSING

3.0

-5 3.5

GO yield cutpoint impact on VDU design. The impact that the HVGO yield has on the VDU column diameter and the VDU fired heater COT is illustrated in Fig. 4. In this example, a FZP of 20 mm Hg(A) is maintained for both the wet and damp VDU designs and, for the wet design, the stripping steam rate is set at 2 wt% on VR product. Fig. 4 indicates that, with the wet VDU design, the required fired-heater COT is reduced by approximately 20°C for a given HVGO yield at the same FZP. However, as discussed earlier, the added volume of stripping steam in the column for the wet design requires a larger column diameter and produces additional sour 30 0

COT (COTMax +/- °C)

Diameter, % DOptimum

108

An additional benefit of steam stripping is that it provides a sharper fractionation between the HVGO and VR products. For the example provided in Fig. 3, the HVGO 95 vol% TBP end point for the wet design with 2 wt% stripping steam is approximately 10°C lower than the damp design. The sharper fractionation is achieved because the HVGO material vaporized with two stages of steam stripping has a lower distillation boiling range than from the HVGO material vaporized with the VDU fired heater. The above effect minimizes the metals content in the HVGO product that improves downstream HVGO conversion unit performance. This may also reduce the nitrogen and carbon residue content in the HVGO product, which may be of benefit to downstream hydroprocessing units. However, the benefits achieved are dependent on the distribution of these contaminates/properties vs. HVGO yield cutpoint as well as the amount of VR entrained into the HVGO draw. For the case of a damp VDU design, the only means of improving fractionation between HVGO and VR products is to increase the wash-oil supply to improve rectification of the flash-zone vapor. This additional wash oil is produced at the expense of a higher VDU fired-heater COT and/or a lower FZP. However, for a deep-cut damp VDU design applied to bitumen-derived AR, a COT increase may not be possible due to the lower COT limitation. The FZP may also be at or near the minimum achievable to meet the minimum overflash requirement. Consequently, additional wash-oil generation may not be possible. Even if a lower FZP is achievable, it will be demonstrated that the VDU column diameter and the motive steam required will increase significantly relative to that of a wet VDU design and is not an economical means of obtaining sharper fractionation.

145 Damp COT

135

-30

125

Wet diameter

Wet COT

-60

115

-90

105 Damp diameter

-120 480

490

500

510

520

530

540

HVGO yield TBP cutpoint, °C FIG. 4

HVGO yield cutpoint impact on VDU design.

95 550

Diameter, % DOptimum

Wet vs. damp VDU design

COT, (COTMax +/- °C)

SPECIALREPORT

CLEAN FUELS

50 FZP

20 Stripping steam

2 0 90

FIG. 5

10 95

100 105 Diameter, % DOptimum

110

diameter due to the COT constraint. Therefore, the wet VDU design is further evaluated. WET VDU COLUMN DESIGN OPTIMIZATION

Optimization of the wet VDU design considers the impacts of the FZP and stripping steam rate on the VDU column diameter and on the motive steam requirements for the overhead ejectors. The objective is to optimize the total steam (motive and stripping) needed and the column diameter based on maintaining an HVGO yield cutpoint of 524°C TBP and the VDU fired heater COT constraint. VDU column diameter. The impact of the flash zone pressure and stripping steam rate on the resulting column diameter is illustrated in Fig. 5. The VDU fired-heater COT is maintained at the recommended maximum and the HVGO yield cutpoint is maintained at 524°C TBP. Fig. 5 indicates that an increase in the stripping steam rate will strip additional material from the flash zone liquid such that the HVGO yield is maintained while permitting reduced vaporization in the flash zone by increasing the FZP. The permitted increase in FZP provides a net decrease in the vapor volumetric flowrate through all sections of the VDU column that reside above the stripping section, even with the additional steam present. The required diameters for these column sections are reduced while that of the stripping section is increased. The reduced diameter of the wash section provides a further benefit in terms of a lower required wash-oil rate to maintain overflash. This benefit is uti-

Flash zone pressure, mm Hg (A)

Stripping steam, wt% of VR

water that must be processed. This example indicates that the damp VDU design with an FZP of 20 mm Hg(A) and the COT constraint, is limited to an HVGO cutpoint of only 507°C TBP resulting in a column diameter that is 6% larger than the determined optimum. The wet design achieves the desired HVGO yield cutpoint of 524°C TBP with a column diameter that is 18% larger than the determined optimum. For comparison, if the pressure for the damp VDU design is reduced to a minimum achievable FZP of, say, 15 mm Hg(A), the estimated HVGO yield cutpoint only increases to 512°C TBP resulting in a VDU column diameter that is 19% larger than the determined optimum. The motive steam required for this case is greater than the total steam required for the optimized wet VDU design resulting in additional sour-water make that must be processed. If a less conservative design basis is established by excluding superheat of the feed, the desired HVGO yield is achievable with a damp VDU design at the COT limit. The resulting FZP for this case is 18 mm Hg(A) and the column diameter is 11% larger than the determined optimum. Another case for comparison is to increase the COT equally for both designs such that the damp design achieves the desired HVGO yield at an FZP of 15 mm Hg(A). For this comparison, a stripping steam rate of 1.4 wt% is selected for the wet design resulting in an FZP of 55 mm Hg(A). The HVGO 95 vol% TBP end points for the above cases are about the same but are slightly greater than the optimized wet design. However, the damp design achieves this fractionation at the expense of an over 40% increase in column diameter, a nearly 50% increase in total steam load and a COT that is 10°C above the recommended maximum. Conversely, the damp design may also be evaluated with applying the determined optimum diameter of the wet design. This damp design case requires an FZP of 28 mm Hg(A) with a corresponding COT that is 16°C above the recommended maximum to achieve the desired HVGO yield. The resulting HVGO 95 vol% TBP end point for this damp design is 7°C greater than the optimized wet design. These examples indicate that the damp VDU design may have difficulties in maintaining the desired HVGO yield within the COT constraint and minimum achievable FZP. Although the wet design requires a stripping section, it does allow for a higher FZP that will decrease column diameter and total steam required. A wet design also provides sharper fractionation in a smaller column

SPECIALREPORT

0 115

Wet VDU column diameter sensitivity. Select 159 at www.HydrocarbonProcessing.com/RS 45

CLEAN FUELS

200

150

FZP

100

10 Stripping steam

0 500

FIG. 6

505

510 515 520 HVGO yield TBP cutpoint, °C

525

450

0 530

HVGO yield cutpoint impact on wet VDU design.

lized to maintain HVGO yield and allow a further reduction of the volumetric vapor rate at the top of the wash section, which allows for a further diameter reduction for the wash and HVGO pumparound sections of the column. Although the wash-oil supply to the wash section is reduced, the degree of fractionation achieved for the cases considered remains relatively constant as the rectification lost in the wash section is negated by shifting vaporization from the flash zone to the stripping section. As indicated in Fig. 5, the described effect is observed until a minimum diameter of approximately 96% of the determined optimum is achieved. As this minimum diameter is approached, the stripping steam rate and the FZP begin to increase exponentially and then will actually reverse to indicate a diameter increase. This minimum occurs at an approximate FZP of 60 mm Hg(A) and a stripping steam rate of 5.3 wt% on the VR product. The minimum diameter is achieved at the maximum efficient application rate of steam (with two stages of stripping) to reduce the column diameter. The permitted increase in FZP becomes insufficient to further reduce the column diameter due to the increasing steam to stripped hydrocarbon lift ratio required to maintain HVGO yield. This means that, to maintain HVGO yield, the corresponding increase in FZP that is permitted due to the incremental HVGO stripped with the additional steam, is insufficient to realize a further decrease in the net vapor volumetric rate profile of the column above the stripping section. GO yield cutpoint. The impact that the desired HVGO yield has

on the flash-zone pressure and stripping steam rate is illustrated in Fig. 6. The VDU fired heater COT is maintained at the recommended maximum with a fixed diameter of the HVGO and wash sections. Fig. 6 illustrates the HVGO yield resulting from the maximum stripping steam rate permitted to maintain the CS factor of a set/existing VDU column for a corresponding/attainable FZP. Similar to the conclusion reached from Fig. 5, but from a revamp perspective, Fig. 6 indicates that more distillate lift is achievable with steam than FZP for a given column diameter, up to a maximum achievable cutpoint. Revamps of existing units should also consider operating at a higher CS factor based on actual unit performance and/or column internal changes that minimize VR entrainment into the HVGO draw. This example is a typical evaluation that may 46

I FEBRUARY 2010 HYDROCARBON PROCESSING

Steam ﬂowrate, thousand bpsd feed

250

Flash zone pressure, mm Hg (A)

Stripping steam, kg/h/thousand bpsd feed

SPECIALREPORT

400 350

VDU sour water

300

Motive steam to 11-PK-101

250 200 150

Stripping steam to 11-C-101

100 50 0 10

FIG. 7

20 30 40 50 Vacuum column ﬂash zone pressure, mm Hg(A)

Sour water production.

be done on an existing facility and would include evaluation of all column sections. Sour-water production. In addition to impacting the VDU column diameter, the FZP impacts both the stripping and vacuum ejectors’ motive steam requirements and, consequently, the sour-water production from the VDU, as illustrated in Fig. 7. For this example, the VDU fired heater COT is maintained at the recommended maximum and the HVGO yield cutpoint is maintained at 524°C TBP. Fig. 7 indicates that an increase in the operating pressure in the VDU column increases the required stripping steam rate but decreases the required motive steam rate. The minimum quantity of sour-water produced (steam consumed) occurs within FZPs of approximately 20 to 35 mm Hg(A). Above this range, any further increase in FZP requires more incremental stripping steam (stripping steam increase illustrated in Fig. 5) than the incremental reduction in motive steam, resulting in a net increase in sour water production. The motive steam is reduced due to higher pressure, moderated with the increase of stripping steam present in the overheads. Similarly, any further decrease in FZP below this range requires more incremental motive steam than the incremental reduction in stripping steam (stripping steam reduction illustrated in Fig. 5), resulting in a net increase in sour water production. The motive steam is increased due to the decrease in pressure, moderated with the decrease of stripping steam present in the overheads. Summary of options. Three main process design parameters associated with recovering distillate material in a VDU are temperature, pressure and steam. For the bitumen-derived AR examples, considerations of these parameters indicate that the damp VDU design may have difficulties in maintaining the desired HVGO yield within the VDU fired heater COT constraint and the minimum achievable FZP. Although the wet design requires a stripping section, it does allow for a higher FZP that will decrease column diameter and the total steam required. A wet design also provides sharper fractionation in a smaller column diameter due to the COT constraint. If the processing objectives are achievable with both design types, selection of a wet vs. damp VDU design is based on the capital and operating expenditures of each option. The largest capital cost impact is the VDU column. The wet design will produce a smaller diameter column but with the addition of a stripping section and the damp VDU design will produce a

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SPECIALREPORT

CLEAN FUELS

Wet designs. If a wet VDU design is selected, the reduced lift required from the VDU fired heater allows for a lower fired heater COT and/or a higher FZP. If the FZP is maintained to minimize the COT, a larger diameter is required for all sections of the VDU column and the stripping steam will increase the sour water production from the VDU. Conversely, if the COT is maintained to maximize FZP, a smaller diameter is required for all sections above the stripping section of the VDU column and the reduction in the motive steam required will decrease the sour water production from the VDU, even with the stripping steam present. For a wet VDU design, the optimum values for the three main process design parameters may be determined with the established payout criteria recognizing the impact that the FZP has on column diameter as per Fig. 5 and that the total steam consumed (sour water produced) is minimized within a substantial range of flash zone pressures considered as per Fig. 7. The stripping section diameter should also be part of the cost evaluation as this section diameter does increase as the ™ ™ ™ Paratherm GLT, HR and MG heat transfer fluids diameters of the sections in the main column give you a host of new and unique benefits. decrease. For the example presented, based on maintaining the coil outlet temperature of the VDU fired heater at the recommended maximum, the FZP selected is 36 mm Hg(A) and the stripping steam rate selected is 2 wt% of vacuum residue product.

larger diameter column without a stripping section. The largest operating cost is that associated with the fuel gas and total steam requirements as well as the cost of treating sour water. Therefore, the largest impact on total cost comparison becomes the desired HVGO yield cutpoint and the VDU fired heater COT constraint imposed due to coking concerns. The COT constraint requires a lower FZP for a desired HVGO yield. For the damp design, this lower FZP requires a substantially larger VDU column diameter and a higher total steam rate. Therefore, deeper and sharper HVGO yield cutpoints coupled with lower COT constraints will tend to drive the optimum toward—and in fact, may necessitate a wet VDU design.

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Processing heavy feeds. In the design or revamp of a VDU that will process bitumenderived AR, success depends on the design concepts and methodology applied. A wet VDU design may not only be optimal but may also be necessary depending on the HVGO yield cutpoint and the VDU fired heater COT limit imposed due to coking concerns. The degree of fractionation required is dictated by downstream HVGO processing requirements and is determined based on the distribution of the contaminants/properties of the feed and the HVGO yield cutpoint desired. The wet design offers sharper fractionation in a smaller VDU column diameter but with the addition of a stripping section. If a wet column design is selected, optimization of operating expenditures (fuel gas and steam consumed, as well as sour water treated, etc.) vs. capital expenditures (column diameter, column stripping section, fired-heater size, sour-water treatment facilities, etc.) can be accomplished with the help of the concepts and methodology presented here. HP

Marco Grande is a principal process engineer with Fluor Canada Ltd., Calgary, Alberta, Canada. His experience is in bitumen/heavy-oil upgrading and downstream refining. Mr. Grande holds a BSc degree in chemical engineering from the University of Alberta.

CLEAN FUELS

SPECIALREPORT

Catalytic technology: Options for better hydrogen production It is not just steam and feed costs that determine processing methods for hydrogen generation K. SCHLÖGL, L. XU, A. DÜKER and W. KALTNER, Süd-Chemie AG, Munich, Germany

ydrogen is and will be one of the most critical utilities within crude oil refineries. This is not just because of the increasing demand of hydrogen for various well-known reasons. Going deeper into the barrel means an ever-increasing demand for hydrogen. The same is true for more heavy and more sour crude oils. As new hydrogen plants must accommodate less equipment to lower capital investment, the design gets tighter and the safety margins in the design are less (Fig. 1). At the same time, longer turnaround schedules and higher production rates are required. This effect is amplified by the need by most refineries to further run at lower margins and to optimize their processes. The costs for all utilities are a high priority, and hydrogen plants are normally evaluated first. Catalysts play a vital role not just in hydrogen plants. In the environment of ever-increasing pressure on the performance, catalysts must be adopted and improved. Catalysts are a determining factor to drive the innovations by technology providers. The technology providers for hydrogen plants take care when optimizing equipment while it is the responsibility of the catalyst supplier, to optimize the operating conditions as determined by the final process flow sheet. The final process design is to a larger extent the optimum operating parameter for the catalysts, which define the technology of the hydrogen plant. Hydrocarbons to hydrogen. Present plant designs show the different steps from hydrocarbons to hydrogen and are evaluated with

a focus on the catalysts with the immediate consequences on the plant design. Special focus is on minimizing steam consumption balanced against maximizing hydrogen production. All possible flow sheets are analyzed, and the benefits from each design are evaluated. The basic decision for the steam-reforming section is whether to implement a pre-reformer or stay with the tubular primary reformer only. The next innovation to be evaluated is shift technology, which is just high-temperature carbon monoxide (CO) shift conversion (HTS) or HTS combined with a low-temperature CO shift conversion (LTS) or a medium-temperature CO shift conversion (MTS). In case of the MTS, the option of an isothermal reactor will be considered. Balancing costs vs. H2 production. We can show that pro-

cess flow schemes depend on many site-specific parameters, such as cost of feedstock as compared to the costs for fuel gas. In the case of cheap hydrocarbon feeds, the efficiency of the total conversion may not be that high, as unconverted hydrocarbons are used as fuel gas to the primary reformer. That combined with the value of steam import against the benefit or penalty for steam export can change the flow sheet of the hydrogen plant completely; it can even overrule the decision that was based on the minimum number of equipment items or optimum amount of export steam. The analyses of several different flow designs are considered with regard to the maximum hydrogen production rate and the lowest steam consumption. Flow schemes with and without prereformer, with and without LTS and with and without MTS will be investigated. The benefits and disadvantages of each configuration will be discussed. In all cases, the feed is a desulfurized natural gas with the same flowrate. The primary reformer is held with a medium heat flux at an outlet temperature of 860°C. The final purification is in all cases done via pressure swing absorption (PSA).

Puriﬁcation

FIG. 1

Hydrogen plant at Neste Oil Oyi (former Fortum) at Porvoo, Finland processing capacity of 155,000 Nm3/h.

FIG. 2

Reformer

HTS

PSA

Common process flow scheme for a hydrogen plant.

HYDROCARBON PROCESSING FEBRUARY 2010

I 49

SPECIALREPORT Puriﬁcation

Reformer

HTS

LTS

PSA

Common process flow scheme for a hydrogen plant plus a low-temperature shift (LTS) reactor.

FIG. 3

Reformer

Puriﬁcation

FIG. 4

CLEAN FUELS

MTS

PSA

Common process flow scheme for a hydrogen plant with medium temperature shift (MTS) reactor in the CO conversion section.

Relative H2 production, %

FIG. 6

Prereformer

Puriﬁcation

FIG. 7

Reformer

HTS

PSA

Common process flow scheme for a hydrogen plant with a pre-reformer and HTS unit.

Prereformer

Reformer

HTS

LTS

PSA

Common process flow scheme for a hydrogen plant with a pre-reformer and HTS and LTS units.

Puriﬁcation

106

Prereformer

Reformer

MTS

PSA

104 102 FIG. 8

100

Common process flow scheme for a hydrogen plant with a pre-reformer and MTS unit.

98 96 Base case-HTS

FIG. 5

HTS-LTS

MTS

Relative hydrogen production – Comparison excluding pre-reforming.

Case histories. Figs. 2–4 represent flow schemes that do not

include a pre-reformer and rely on the primary reformer. Fig. 2 shows one of the most common flow schemes for a natural gasbased hydrogen plant. It comprises a primary reformer, an HTS reactor and a PSA. The flow sheet in Fig. 3 adds an LTS reactor downstream of the HTS unit. The CO conversion section in Fig. 4 comprises an MTS reactor. The first task is to determine the minimum steam required to the process gas. The limitation for steam in all three cases is the minimum steam/carbon ratio (S/C ratio) at the inlet of the primary reformer. The S/C ration depends on several parameters—including composition of the natural gas feed, average heat flux of the furnace, operating pressure outlet temperature—and, of course, on the catalysts. Using advanced steam reforming catalysts, the S/C ratio can be as low as 2.6 to 2.8 mole/mole. If the S/C ratio drops below this value, carbon will form in the reformer tubes leading to negative consequences that are well known to operators of steam reformers. Limits of the S/C ratio at the reformer inlet are also inline with the requirement of a minimum amount of steam to the HTS catalyst. All commercially available HTS catalysts are Fe-CrOxides catalysts that over-reduce when the steam/hydrogen ratio drops below certain limits. The consequence of over-reduction 50

Puriﬁcation

I FEBRUARY 2010 HYDROCARBON PROCESSING

is the formation of higher hydrocarbons and oxygenates by the Fischer-Tropsch (FT) reaction. Improved HTS catalysts mitigate byproducts formation via over-reduction at S/C ratios as low as 2.6 mole/mole at the reformer inlet. It is shown later that the steam can further be reduced without any negative effect on the performance of the HTS catalyst. As in our study, the S/C ratio, the feed flowrate and reformer outlet temperatures are fixed; hydrogen production now depends only on the outlet temperature of the CO conversion reactor. The lower the temperature, the more beneficial the equilibrium towards lower CO concentrations; thus, more hydrogen is produced. Fig. 5 shows the results of the calculations. The outlet temperature of the LTS reactor approximately 80°C lower than the MTS. This is again approximately 100°C lower than the outlet temperature of the HTS. Consequently, the hydrogen production is highest when there is an LTS reactor downstream of the HTS. To further reduce the steam needs, a pre-reformer can be added to the reforming section. This will lower steam requirements at the primary reformer inlet, as the pre-reformed gas is much easier to handle in the fired-tubular reformer. The limitation for the lowest steam consumption is not the reforming section but the catalyst for the CO conversion. Even the best HTS catalyst cannot operate at the low steam levels that are now possible with the addition of a pre-reformer. This disadvantage is limited to iron (Fe) based catalysts only. The catalysts for LTS and MTS CO conversion are copper (Cu) based and do not have the limitation on steam requirement. This benefit, however, can only be captured when there is no HTS unit. The LTS catalyst downstream of the HTS reactor may be able to run at low steam. However, it is the HTS

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SPECIALREPORT

CLEAN FUELS

Relative steam consumption, %

100

60 Base case HTS-LTS HTS FIG. 10

MTS

PREREF PREREF Base case HTS-LTS HTS

PREREF MTS

Relative steam consumption—all cases.

105 Top-fired primary reformer.

■ The optimum operating parameters

of catalysts define the technology of the

100

Relative H2 production, %

FIG. 9

hydrogen plant.

95 90 85 80 75

I FEBRUARY 2010 HYDROCARBON PROCESSING

PREREF Base case HTS FIG. 11

PREREF HTS-LTS

PREREF MTS

Relative hydrogen production—comparison including prereforming.

105 100

Relative H2 production, %

catalyst that determines the higher S/C ratio. So, installing the LTS will not influence process steam requirements. The flow sheets in Figs. 6–8 are comparable to the ones in Figs. 2–4 with the difference of the pre-reforming section. The steam to the hydrogen plant is now less as the pre-reformer can run at an S/C ratio with natural-gas feed as low as 0.5 mole/mole. Fig. 10 shows the results of all studies. The steam requirement for the cases with the HTS unit is 25% less and, in the case of the MTS, the steam requirement is even up to 40% less compared with the cases running without the pre-reformer. The extreme low-steam values are only achieved by splitting the steam to the individual units. Part is fed to the pre-reformer and the balance is added upstream of the primary reformer. Thus, the overall S/C ratio of 1.6 mole/mole can be achieved at the primary reformer inlet. This low S/C ratio is only possible with the Cu-based MTS catalyst downstream. In Figs. 6 and 7, where the HTS catalyst limits the steam requirement, another steam addition to the HTS unit is required. This additional step is the steam quench to the HTS unit; it adds complexity to the plant and is normally not an option. Therefore, the balance of the steam is fed to the primary reformer to achieve an overall S/C ratio at the primary reformer inlet of 1.8 mole/mole. With lower steam flow to the hydrocarbon feed, the overall conversion rate to hydrogen is, of course, less. Fig. 11 shows that, in designs with pre-reforming, the hydrogen production is approximately 20% less as compared to the cases with higher steam addition. It is still the outlet temperature of the CO conversion reactor that controls hydrogen production rate; the lower the temperature, the higher the hydrogen production. Fig. 12 summarizes the results for all cases. It was the task to find the design

95 90 85 80 75 Base case HTS-LTS HTS

FIG. 12

MTS

PREREF PREREF Base case HTS-LTS HTS

PREREF MTS

Relative hydrogen production—all cases.

with the absolute minimum steam requirement. This is the case where a pre-reformer and an MTS are installed. When considering steam and feed consumption as the only criteria, the decision for a technology can be easy. The cheaper the hydrocarbon feed, the less efficient the process has to be, means the less hydrogen will be produced at a given feedrate. A certain hydrogen production rate requires approximately 20% more

CLEAN FUELS hydrocarbon feed in the case of a flow sheet with the MTS unit as compared to the standard flow sheet comprising reference-HTS only. At the same time, approximately 40% less steam is required. Each site has its specific costs for steam and hydrocarbons; therefore, each site will come up with different results. It is not just steam and feed costs that must be considered in the decision-making for hydrogen technology. The catalysts need certain optimum conditions and special startup and shutdown procedures, which can influence the decision for a certain technology. Options. Many factors influence the decision-making over

service for catalysts for the production of ammonia and methanol, technologies for crude oil refineries as well as novel solutions for the new generation of alternative fuels via Fischer-Tropsch synthesis. Dr. Düker holds a PhD in chemistry from the University of Munich, Germany.

Wolfgang Kaltner is member of the sales team for catalytic technologies for the refinery industry in the headquarters of Süd-Chemie AG in Munich. He works closely with engineering companies in the design of new hydrogen plants. Dr. Kaltner studied chemical engineering and holds a PhD from the University of Munich, Germany.

hydrogen technology. The costs for hydrocarbon feed and the value of steam consumption and steam export must be considered as well as investment costs, complexity and reliability of operation and flexibility in the feedstock. The absolute lowest consumption of steam can be achieved by a process design comprising pre-reformer, primary reformer and an MTS. This is only possible, however, on cost of hydrogen yield and only recommended when the hydrocarbon feed is almost as cheap as fuel gas. The highest hydrogen production is achieved by the simple flow sheet comprisExtensive ing a primary reformer and an HTS. The SENTRON higher steam consumption can be compensated by a higher steam export. The solutions of all these questions are given very early in any project by the close cooperation of the hydrogen plant technology provider with the project owner and supplier of catalytic technology. HP

field experience demonstrates that delivers Tangible Savings Solutions.

Kerstin Schlögl is a product manager for catalytic technologies for Süd-Chemie AG in Munich. She has several years of experience in catalysis in oil refinery and coal-based industries. Her technical experience includes R&D, sales and technical service. Dr. Schlögl holds a PhD in chemistry from the University of Ulm, Germany.

“ Ling Xu is a product manager for catalytic technologies for Süd-Chemie AG in Munich. She has several years of experience in catalysis for oil refinery and coal-based industries. Dr. Xu has worked in many areas including R&D, technical service and sales and holds a PhD in chemistry from the University of Iowa.

Axel Düker is the director of sales for catalytic technologies for the refinery industry in the headquarters of Süd-Chemie AG in Munich. He has over 18 years of experience in catalysis and his professional career covered sales and technical

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CLEAN FUELS

SPECIALREPORT

Dealing with dieselization Several processing options can provide cost-effective ways to maximize diesel make and limit gasoline production

Crude and cutpoint selection. The crude processed in the

refinery and the cutpoints used to fractionate that crude can have a significant impact on the diesel production from the refinery. Diesel is produced from two main sources in crude straight-run (SR) materials around 150°C–360°C and diesel produced in the upgrading of vacuum gasoil (VGO) or residues. Fig. 3 illustrates the impacts from different crudes on diesel yields. This figure shows the yields of 150°C–360°C SR distillate material and the yields of diesel from 360°C–560°C VGO material assuming a 30% yield conversion yield.

1990 1992 1994 1996 1998 2000 2002 2004 2006 2008 2010 2012 2014 2016 2018 2020 2022 2024 FIG. 1

Road diesel

Non-road diesel

European diesel and gasoline markets. (Source: Wood Mackenzie—Q1 2009)

4.0 3.5 3.0

Total diesel to gasoline

2.5 2.0 1.5 1.0

Road diesel to gasoline

0.5 0.0 1990 1992 1994 1996 1998 2000 2002 2004 2006 2008 2010 2012 2014 2016 2018 2020 2022 2024

Gasoline and diesel markets. Fig. 1 shows the projected markets for gasoline and diesel in Europe from 1990 to 2025 as predicted by Wood Mackenzie. To illustrate what this means for refiners, Fig. 2 shows the diesel-to-gasoline ratio in Europe. Options for improving diesel yields. Several options are available to improve diesel make (or reduce gasoline make) from a refinery, and, in particular, for improving the diesel-to-gasoline ratio produced from the refinery. These options include: • Crude and cutpoint selection • Improved diesel recovery • Targeting diesel production on the FCC • Residue upgrading options • Converting LPG and naphtha to diesel • Converting gasoline into petrochemicals. We will consider each option and investigate the impacts that these options can have on the refinery balance and potential to achieve a 100% diesel refinery.

Diesel and gasoline demands in Europe

450 400 350 300 250 200 150 100 50 0

Gasoline

Europe diesel to gasoline ratio

efiners face a significant challenge over the coming years as global transport fuels see a continuing increase in demand for diesel at the expense of gasoline. There are several proven potential options available to improve diesel make from a refinery. We will investigate a number of scenarios to demonstrate the potential benefits of available processing schemes to increase the diesel to gasoline ratio. Key consideration areas include crude selection, optimizing cutpoints, improving diesel recovery from the vacuum distillation unit (VDU,) modifying fluid catalytic cracking (FCC) yields, installing residue upgrading, technologies for converting liquefied petroleum gas (LPG) and gasoline to diesel and converting gasoline to petrochemicals. Several examples will investigate the advantages and disadvantages of each scheme and highlight some potential synergies among available options, as well as include capital and operating cost impacts, impact on overall refinery yield and consideration of carbon emissions.

Diesel and gasoline demand, million tons/yr

M. STOCKLE and T. KNIGHT, Foster Wheeler, Reading, Berkshire, UK

FIG. 2

Diesel to gasoline ratio in Europe. (Source: Wood Mackenzie—Q1 2009)

TABLE 1. Impact of cutpoints on gasoline and diesel yields Case

Max. gasoline wt% yields on feed

Max. diesel wt% yields on feed

LPG

2.90

2.65

Gasoline

38.73

37.02

Kerosine

Diesel

28.78

31.81

Fuel oil

23.20

22.31

Diesel: Gasoline

0.75

0.86

HYDROCARBON PROCESSING FEBRUARY 2010

I 55

SPECIALREPORT

CLEAN FUELS TABLE 3. Impact of residue upgrading schemes on diesel make

Potential diesel yields from crude

70 VGO Diesel

60 50

Case

40 30 20

HCK + VIS wt% yield on feed

FIG. 3

WTI

Arab Tapis Bonny light light

Boiling points of hydrocarbons.

TABLE 2. Impact of gasoline yields Case

FCC max. gasoline wt% yields on feed

FCC max. diesel wt% yields on feed

LPG

2.90

6.45

Gasoline

38.73

27.19

Kerosine

Diesel

28.78

37.73

Fuel oil

23.20

22.37

Diesel: Gasoline

0.75

1.39

Maximizing diesel yields by selecting the right cutpoints within the refinery is also important. Table 1 lists the overall refinery mass balance for an FCC refinery where the only change is modifying cutpoints in key distillation units while prices are set to maximize gasoline in one case and diesel in the other. Improve diesel recovery. For many refineries, the back-end boiling point specification limits the amount of diesel that can be produced. Since diesel is the heaviest cut produced in the crude unit, it can be difficult to achieve good separation between the diesel and residue cut. Carryover of residue material into the diesel cut can limit recovery of diesel-range material and reduce diesel yield. The nature of crude unit design means that it is difficult to get good separation at the bottom of the crude unit; by allowing some diesel to slip into the vacuum unit and recovering a light vacuum gasoil (LVGO), good separation can be achieved between the LVGO and VGO as it is at the top of the column and more diesel-range material can be recovered before the boiling point limitation is reached. This could result in 1%–2% of crude being recovered as diesel rather than being routed to the VGO upgrading unit. The actual impact of this on the petrol-to-diesel ratio will depend on the upgrading units installed and the separation quality achieved in the refinery. Improved diesel production from the FCC unit. The

focus by many European refiners has been on improving the yields of diesel-range material from their VGO upgrading unit. For many, this has meant maximizing light cycle oil (LCO) make. This has brought its own challenges as LCO is not a particularly good diesel blendstock, often having a low cetane and high density.

I FEBRUARY 2010 HYDROCARBON PROCESSING

FCC + RHCK wt% yields on feed

2.90

3.06

3.09

Gasoline

38.73

19.25

39.15

39.90

Kerosine

5.13

Diesel

28.78

47.76

43.91

36.08

Fuel oil

23.20

22.34

13.03

0.75

2.5

Diesel: Gasoline Maya Arab Brent Flotta Urals heavy

FCC + COK wt% yields on feed

LPG

Coke

10 0

FCC + VIS max. gasoline wt% yields on feed

6.68 1.12

0.90

However, for many refiners, LCO has been about making the most of the assets they have. Recent years have seen particular focus on this area from all FCC licensors and, in particular, the quest for the ideal combination of high propylene and high LCO, which have tended to be mutually exclusive in conventional FCC units. This subject has been covered in many publications from FCC licensors and will not be discussed here in detail. However, to illustrate the possible impacts of optimizing FCC yields, Table 3 shows the overall mass balance for a conventional FCC- based refinery in which the FCC operation is maximizing gasoline and distillate make. Residue upgrading options. One option to improve diesel

yields is to seek additional residue upgrading capacity within the refinery. There are a number of options, including: • FCC • Hydrocracker • Residue hydrocracking • Coking. Table 3 compares overall product slates from combinations of some of these residue upgrading options. For many refiners, the main VGO upgrading unit is already fixed, and the question becomes how other upgrading options can be combined with an FCC unit. From Table 3, combining FCC with vacuum residue conversion can improve diesel make and the diesel to gasoline ratio, with both coking and residue hydrocracking showing improvements in the diesel make. Another option is to add VGO hydrotreating to the configuration. If the VGO hydrotreater (HDT) is run as a mild hydrocracker, we can produce some diesel from the unit and also improve quality of the FCC products. Of course, we face the drawback of converting some FCC feed and must run the unit at turndown. There are options to increase VGO make from the refinery such as adding a coker or solvent deasphalting (SDA) unit and carefully balancing the additional VGO make and conversion of the VGO HDT can allow a refiner to add upgrading while making the most of the existing facilities. Table 4 shows the yields of a typical FCC refinery with the addition of a VGO HDT and SDA/coking unit. In this case, the SDA pitch is assumed to be used for onsite power generation. We can see from Table 4 that adding an SDA/coking unit can further increase the diesel yield of the refinery. The options considered prior to this section have required little capital investment and have had only relatively small impacts on operating costs. Investing in new upgrading units is a different story with significant investment costs being required and a significant increase in operating costs can also be expected

CLEAN FUELS TABLE 4. Adding upgrading to an FCC refinery

BP of typical hydrocarbons

450 LPG

3.72

Gasoline

38.83

Kerosine

Diesel

40.72

Fuel oil

3.38

Diesel: Gasoline

1.05

Converting LPG and gasoline to diesel. So far, the

options we have considered used materials already in the diesel range or cracked materials that are too heavy for processing into lighter diesel. There is only so much that can be achieved; we will look at converting lighter material into diesel. Fig. 4 shows the boiling range of some typical hydrocarbons and highlights the boiling range of potential diesel material. We are looking for molecules with 11 to 20 carbon atoms to produce diesel. Also, we need to look for technologies that convert C3–C10 material to C11+ . A number of potential paths are already in use in the refinery and petrochemicals industry for converting LPG and aromatics to heavier hydrocarbons. The options revolve around two main technologies, olefin oligomerization/polymerization and benzene/ light aromatic alkylation. Oligomerization of C3 and C4 olefins to produce a C6 or C8 gasoline blendstock is a well-established technology with at least two licensors offering proven technology. This technology is

BP of typical hydrocarbons, °C

FCC with SDA/coking unit, wt% yields on feed

SPECIALREPORT

N-paraffin I-paraffin Naphthene Mono-aromatic Poly-aromatic Olefin

400 350 300 250 200 150 100 50 0 5

FIG. 4

13 15 17 Carbon number

Potential diesel yields from crudes.

not just limited to C3 and C4 feeds, and C5 –C8 material can be recycled back to the unit to produce higher carbon number products in the kerosine range. The material is olefinic and requires hydrotreating before it can be blended into the diesel pool. Once processed, it is a high-cetane, low-density blendstock. Table 5 lists the impact of converting an alkylation unit into an oligomerization unit to produce a kerosine and gasoline product. One of the key limitations of this technology is that it can only convert olefins; saturated material passes through unconverted. There are technologies available (and commercially proven) to produce olefins from saturates, and combining these with

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HYDROCARBON PROCESSING FEBRUARY 2010

I 57

SPECIALREPORT

CLEAN FUELS

oligomerization offers the opportunity to significantly increase LPG conversion and increase distillate yields. We have already seen that we can recycle the C5–C8 material produced in the oligomerization unit to produce kerosine and, therefore, it should also be possible to feed C5–C8 olefins produced elsewhere in the refinery into the unit to increase yields. This does create some problems, since oligomerization technologies are sensitive to impurities present in these materials. Removing impurities to a sufficient level also tends to saturate the olefins. One option could be to remove the impurities and then recreate the olefins, thus producing olefins from SR material using dehydrogenation technologies. This offers the potential to increase distillate yields further, although the technology is more challenging when applied to heavier hydrocarbons and it is expensive. Figs. 5–7 show several options for combining oligomerization and dehydrogenation into FCC-based refinery flow schemes. TABLE 5. Alkylation vs. oligomerization FCC and alkylation wt% yields on feed

FCC and oligomerization wt% yields on feed

LPG

2.90

4.41

Gasoline

38.73

33.04

Kerosene

Diesel

28.78

34.14

Fuel oil

23.20

22.20

Diesel: Gasoline

0.75

1.03

A second option for converting LPG range material to kerosine-range material is aromatic alkylation technologies. There are several technologies based on the alkylation of benzene (and toluene or heavier) with olefins that offer the potential to produce distillate-range material from LPG and gasoline-range material. Technologies that could be suitable include those used for benzene reduction to produce linear alkyl benzene (LAB) and those used for cumene production. While these processes are used for petrochemicals production, they need to be looked at in a different way and may require “detuning” to maximize production of distillate-range material rather than high-purity petrochemicals. Figs. 8 and 9 illustrate a conventional cumene production process and how this could be detuned to maximize diesel production rather than targeting petrochemicals. The key potential changes are: • Broaden the range of the olefin feed. Use mixed LPG and possibly even light naphtha-range olefins or mixed olefins and paraffins; use mixed or lower purity aromatics, including toluene as well as benzene. • Accept or even increase byproducts and remove conversion of byproducts. Remove conversion of diisopropylbenzene (DIPB). • Change reactant ratios to produce the best diesel-range material. For example, increase ratios of olefins to aromatics to increase the amount of DIPB produced. Hydrogen Light naphtha

Dehydrogenation

Saturated LPG

Dehydrogenation

Saturated LPG

FRAC Mixed LPG

FRAC

Polymate Mixed LPG

Oligomerization

Polymate

Oligomerization

Kerosine

Kerosine FIG. 5

FIG. 7

Oligomerization with recycle.

Benzene

Hydrogen Saturated LPG

Dehydrogenation

FRAC C3=

FRAC Mixed LPG

Cumene

Alkylation reactor

Polymate

Oligomerization

FRAC

Transalky reactor

Kerosine FIG. 6

Oligomerization with LPG and light naphtha dehydrogenation.

Oligomerization and LPG dehydrogenation.

I FEBRUARY 2010 HYDROCARBON PROCESSING

FIG. 8

Conventional cumene process.

DIPB

CLEAN FUELS

SPECIALREPORT

Saturated LPG

Dehydrogenation LPG

LPG

FRAC

NHT

CCR

CDU

Aromatic extract

Aromatic alky

Diesel

DHT

Mixed aromatics FRAC Mixed LPG

HCK

Alkylation reactor

DHT

VDU FCC Coker

Heavy aromatics FIG. 9

If these processes can be modified to process C5â&#x20AC;&#x201C;C8 olefins rather than LPG as a feed, then some interesting combinations with oligomerization become possible. For example, light naphtha-range material and saturated LPG can be fed to dehydrogenation units; the product is combined with olefinic LPG and fed to an oligomerization unit. The distillate product from this is routed to the hydrotreater, unreacted saturated LPG is recycled to the dehydrogenation unit and the C5â&#x20AC;&#x201C;C8 material is partially recycled and partially routed to a benzene alkylation unit where it is reacted with benzene, toluene and possibly heavier aromatics extracted from the reformate. If we combine this with changes in

LPG

Coke to fuel

Detuned cumene process.

LPG

FIG. 10

Dehyd

Olig

Zero gasoline refinery.

the operation of the FCC, we may be able to move to zero gasoline from an FCC-based refinery, as shown in Fig. 10. This level of gasoline and LPG conversion is not commercially proven and would undoubtedly require significant capital investment, but it does demonstrate that there are some interesting options put forward by these technologies. It is also interesting to note that in this configuration, although we no longer make gasoline, we will still have hydrogen production from the CCR and can add hydrogen from the dehydrogenation unit.

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SPECIALREPORT

CLEAN FUELS

TABLE 6. Comparison of options for improving diesel production Case

Base FCC

High LCO

Oligomerization

FCC & Coker

0.75

1.39

1.03

1.12

FCC and Ascot

HCK

FCC and RHCK

1.05

2.5

0.90

Diesel: Gasoline Case Diesel: Gasoline

rial to feed either a BTX production facility or as a feedstock to a steam cracker for ethylene and propylene production. These options require significant investment and diversification of products produced from the refinery. But they can offer good returns under the right circumstances, especially if the right point in the petrochemicals cycle is chosen for the investment. Comparison of results. Table 6 summarizes the gasoline-

TABLE 7. Impact of changes on other aspects of the refinery Aromatics Cutpoints Recovery FCC yields Oligomer alkylation Capital impact Opex impact

Low

High

Minimal

Low

Medium

TABLE 8. Impact of changes on CO2 emissions for FCC-based options FCC + FCC + VIS max. VIS max. gasoline diesel Relative CO2 emissions

1.07

FCC + COKER

FCC + SDA/ coking

FCC + RHCK

Oligomer

1.11

1.22

1.19

Converting gasoline to aromatics or petrochemicals.

One final way to change the gasoline the diesel ratio from the refinery is not to increase the diesel yield but to decrease gasoline make. The obvious option here is to use the naphtha-range mate-

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to-diesel ratio produced from the various refinery configurations we have considered. Each of these represents one step away from the base case and, by combining a number of these options, more substantial improvements in diesel production can be achieved. Table 7 shows a comparison of some of the other impacts that making the changes discussed in this article will have on the refinery. The carbon emissions of each scheme are also an important consideration and Table 8 lists the relative emissions of some of the presented options. We have seen that there are a number of options available to refiners. Many refiners will already have pursued some of these options, but virtually all refiners will still have some options available to them. The key is picking the right options for the individual position a refiner is in. To ensure that a refiner selects the best options available to them, they will need to consider these factors: â&#x20AC;˘ Ensure that the costs of any scheme are realistic; obtain feedback from recent projects to ensure that any estimate reflects current changes in costs and the individual requirements of the market in the project location. â&#x20AC;˘ Verify that any advice received is provided by a company that understands the many options and technologies involved, but are technology independent and unbiased in its opinion. â&#x20AC;˘ Consider the impact of any changes on other refinery systems such as utilities and hydrogen. â&#x20AC;˘ Investigate the impact of any changes on carbon emissions. If these factors are evaluated and if schemes are selected that balance the competing requirements listed above to provide the most economic solutions, then refiners can increase their diesel yields significantly and could even develop a refinery that makes 100% diesel. HP ACKNOWLEDGMENT Revised and updated from an earlier presentation at the ERTC 14 Annual meeting, Nov. 9â&#x20AC;&#x201C;11, 2009, Berlin, Germany.

Mike Stockle is chief engineerâ&#x20AC;&#x201D;Refining Technology and currently works in Foster Wheelerâ&#x20AC;&#x2122;s Business Solutions Group in Reading, UK. He graduated from Nottingham University in 1995 and is a Chartered Engineer and a Full Member of the IChemE. During his time at Foster Wheeler, he has worked on a number of refining projects ranging from a grassroots refinery configuration studies and FEEDs, through major refinery revamps, to a number of small studies. Mr. Stockle is an experienced LP modeller and has undertaken a number of feasibility and front-end studies looking at the impacts of changing markets and legislation on refineries across the globe.

Tina Knight is a senior process engineer within Foster Wheelerâ&#x20AC;&#x2122;s Business Solutions Group in Reading, UK. She has worked for Foster Wheeler since graduation in a number of process engineering roles including a number of refining projects ranging from conceptual studies through to EPC. Ms. Knight has developed linear programming capability over the past few years and has used it to complete several grassroots refinery configuration studies. She graduated from Loughborough University in 2003 with a MS degree in chemical engineering and has been a Chartered Engineer and Full Member of the IChemE since 2008.

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CLEAN FUELS

SPECIALREPORT

Improve usage of regenerated refining catalysts Common-sense guidelines detail when it is beneficial to reuse catalysts G. J. YEH, Saudi Aramco, Dhahran, Saudi Arabia

arious catalyst types are used in the refinery catalytic processesâ&#x20AC;&#x201D; hydrotreating, hydrocracking, isomerizaton, fluidized catalytic cracking (FCC), naphtha reforming, alkylation, hydrogen generation, etc. Catalysts play a major role in petroleum refining. Catalysts accelerate reactions, so that desired reactions can occur at lower temperatures and pressures than those required for noncatalytic processes. In addition, catalysts do deactivate during the process, and must be regenerated or replaced at the end of the cycle. Annually, refiners worldwide spend billions of dollars to purchase refining process catalysts. Some spent catalysts can be regenerated to recover their activities, and reused in the same processing units or in less-severe services. If the regenerated catalyst can meet performance requirements, using regenerated process catalyst can improve refinery operating margins. It is a common practice to use the regenerated catalysts in crude oil refineries. A major Middle East oil company uses regenerated hydrocracking catalysts up to three regenerations by ex-situ and regenerated semi-regen naphtha reforming catalysts up to 10 regenerations by in-situ. Utilization of the regenerated hydrotreating catalysts has been recommended in less-severe applications, i.e., naphtha and kerosine TABLE 1. Critical levels of typical metal contaminants Contaminants

Critical level, wt%

Sodium

0.2

Silicon

Ni + V

Arsenic

0.1

Iron

hydrotreaters, but has not been implemented in the company. We will explore common practices in the refining industry regarding application of regenerated hydrotreating, hydrocracking and semiregen naphtha reforming catalysts. Presented guidelines over regenerated catalysts show how refiners can save by using regenerated catalysts or enhance revenue by using the fresh catalysts. HYDROTREATING CATALYSTS

Various hydrotreater designs are used in oil refineriesâ&#x20AC;&#x201D;light straight-run naphtha hydrotreaters, naphtha hydrotreaters, FCC gasoline post-treaters, kerosine hydrotreaters, diesel hydrotreaters, vacuum gasoil (VGO) hydrotreaters, atmospheric residue desulfurizers (ARDs) and vacuum residue desulfurizers (VRDs). Hydrotreating catalysts used in these hydrotreaters are CoMo and NiMo, and very occasionally, NiW supported on gamma alumina. Hydrotreating catalysts must be sulfided to be active in removing metals, sulfur and nitrogen, and in saturation of olefins and aromatics. During the process, hydrotreating catalysts can be irreversibly deactivated by coke formation, metal poisoning, reduction of active metals and sintering. At the end of the cycle, the spent hydrotreating catalyst could be regenerated to recover activity and reused in the same unit or less-severe services. As a rule of thumb, the regenerated hydrotreating catalysts should have at least 80%â&#x20AC;&#x201C;85% of the fresh catalyst surface area and have a low-level of metal contamination to be reused. Spent hydrotreating catalysts containing high metal contents cannot be regenerated and reused, such as the hydrotreating catalysts used in ARDs, VRDs, hydrocracker pretreaters and VGO

hydrotreaters, and naphtha and GO hydrotreaters treating coker feedstocks. Metal contaminants remain on the hydrotreating catalyst after regeneration, resulting in low activity and surface area recovery. To use the regenerated hydrotreating catalysts with metal contaminants, the top layer of the catalyst bed must be vacuumed and discarded. Table 1 summarizes common metal contaminants and their critical levels. Catalysts containing any metal contaminant higher than critical levels are generally deemed nonregenerable, and it is not economical to reuse them. However, it is always better to contact the catalyst supplier regarding critical levels of metal contamination for their catalysts. In most hydrotreaters, a guard bed is used at the top of the reactor to mitigate pressure drop buildup and to protect the main hydrotreating catalyst from contamination. It is never economical to regenerate the guard-bed catalyst. Hydrotreating catalysts with significant surface area loss are not regenerable. Catalysts used in light straight-run naphtha hydrotreaters, naphtha hydrotreaters, FCC gasoline post-treaters, kerosine hydrotreaters, diesel hydrotreaters and VGO hydrotreaters, can be regenerated and reused. To justify using regenerated hydrotreating catalysts, the regenerated catalyst should have good crush strength, particle length, and sufficient activity and stability to provide a reasonable cycle life. Industrial practices. The

hydrotreating catalysts are typically regenerated once or twice and reused in the same units or in less-severe services if they are suitable. Most regenerations are conducted ex-situ, since ex-situ regenerations provide a better regeneration than in-situ. HYDROCARBON PROCESSING FEBRUARY 2010

I 63

SPECIALREPORT

CLEAN FUELS

Some refineries regenerate hydrotreating catalysts in-situ if there are no regeneration facilities available to them. For in-situ regeneration, it is impossible to monitor regeneration temperature and get a proper regeneration. Consequently, sintering and metal agglomeration can occur, resulting in significant activity loss. The regenerated catalysts can be reused in the same unit or can be cascaded to less-severe services. It is a good synergy to use the fresh hydrotreating catalysts in severe ser-

vices, e.g., ultra-low-sulfur diesel (ULSD) units, etc., and the regenerated catalyst in less-severe applications, e.g., naphtha hydrotreaters, kerosine hydrotreater and diesel hydrotreaters, etc. Regen hydrotreater cat. If the hydrotreating catalyst is properly regenerated, the effects of each regeneration ex-situ is listed in Table 2. The new generations of hydrotreating catalysts with Type II active sites need an additional rejuvenation step

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besides coke burning to recover most of the activity up to 90%; normal regeneration steps will only recover about 70% of the new generation hydrotreating catalyst activity. This additional step can be costly, and the only way to decide whether to proceed with the additional rejuvenation step is to do an economic analysis. Economic justification. Whether to

use the regenerated or the fresh hydrotreating catalyst is an economic decision. The main factors to consider include: • Cost of regeneration • Cost of fresh catalyst • Impact of a shorter cycle on maintenance and production cost • Benefits of new catalysts. When regenerated catalyst is used in a less-severe application, the economic justification is most likely favorable as long as the regenerated catalysts can meet the required performance. If the regenerated catalyst is used in the same unit, the economic impact of a shorter cycle and possible benefits of a new and better catalyst should be considered. Hydrotreater example. A refin-

Fiberflex® for Difficult Measurements High viscosity, corrosive characteristics, and varying specific gravity makes refining’s measurement of heavy crudes difficult. Ohmart/VEGA’s FiberFlex radiation-based level measurement system, is up to the task. This non-intrusive system is unaffected by extreme process characteristics. The detector’s flexible sensor is able to follow the contour of rounded and conical vessels which make it ideal for fractionators and vacuum bottom measurements. Advantages: t Unaffected by high process temperature and pressure t Longest detector available in the industry (up to 23 feet) t Flexible to fit vessel geometry

t Lightweight for easy mounting t Offers a wide variety of setup and compensation options t HART, Foundation Fieldbus, and Relay outputs

ery had about 70,000 kg of spent diesel hydrotreating catalyst in the warehouse. This catalyst batch has only been used in the diesel hydrotreater for five years and has never been regenerated and/or reused in other units. Analytical results show that the regenerated catalyst recovers 85%–90% of surface area of the fresh catalyst and possesses good crush strength, and the metal contaminants on the regenerated catalyst are low. This batch of catalyst can be regenerated and reused in naphtha and kerosine hydrotreating to meet the unit operation requirements. Therefore, it was recommended accordingly. The economic benefit to use this batch of the regenerated catalyst was determined. The breakdown costs for the fresh catalyst purchase and the regenerated catalyst are summarized in Tables 3 and 4, respectively. TABLE 2. Effect of each regeneration on hydrotreating catalyst properties and performance Factor

Select 163 at www.HydrocarbonProcessing.com/RS 64

~10

Activity

~10

Cycle length

www.ohmartvega.com info@ohmartvega.com 800.FOR.LEVEL

% loss

Surface area (SA)

~10

Crush strength

negligible

Product yield

negligible

TABLE 3. Cost for fresh catalyst purchase1 Description

Cost, $/kg

Catalyst cost

TABLE 5. Effect of regeneration on the hydrocracking catalyst properties and performance

19.14

Factor

% loss 15–301

0.35

Surface area

Potential sale benefit3

-0.40

Activity

Total

19.09

Cycle length

Transportation

1 2 3

cost2

Crush strength

Based on 70,000 kg of catalyst. From a typical manufacture site in Europe or US West Coast. If sold to metal reclaimer.

TABLE 4. Cost for regenerated catalyst1 Description

Cost, $/kg

10 10–15 Negligible

Product yield

1–2 vol% for middle distillate 0.5–1 vol% for C5+ liquid

Product quality

Jet smoke point (1–2 mm), diesel cetane (1–2 numbers)

The SA loss on zeolitic hydrocracking catalyst after regeneration can be as high as 30%.

Catalyst regeneration cost

1.60

Industrial practices. Hydrocracker

Inert stripping

0.4

Screening

0.15

Length grading

0.30

pretreating catalysts are typically regenerated one time or never. The top of the pretreating catalysts may be skimmed and then discarded due to high-metal contents if the pretreating catalyst is to be regenerated. The activity of the pretreating catalyst is normally the bottleneck in the hydrocracker. Regenerated pretreating catalyst is usually used in a less-severe application instead of returning to service in the hydrocracker. New generations of hydrocracker pretreating catalysts with Type II active sites require an additional rejuvenation step to recover 90% activity. This additional step can be expensive. The only way to decide whether to proceed with this additional rejuvenation step is to perform an economic analysis. Several refiners regenerate their hydrocracking catalyst only once. Few regenerate the hydrocracking catalyst more than once. Most refineries don’t even regenerate hydrocracking catalyst and purchase fresh catalyst each time due to economic reasons—cost due to reduced performance is higher than the cost of the new fresh catalysts. Compared with the zeolitic hydrocracking catalysts, the amorphous hydrocracking catalysts have a lower activity and a higher selectivity for distillate products. When regenerating hydrocracking catalyst occurs, acid sites are destroyed. Generally, the zeolitic hydrocracking catalysts have a much higher number of acid sites than amorphous type catalysts. Therefore, the zeolitic hydrocracking catalysts suffer a lower activity loss than the amorphous hydrocracking catalyst during regeneration. If the coke laydown is extremely high, as is common for high-severity units, the burn temperature during regeneration can be difficult to control. This can lead to met-

Transportation cost

0.12

Total

2.57

Based on 70,000 kg of catalyst.

Assuming a loss of 5% during the regeneration process, the total savings to use this batch of regenerated hydrotreating catalyst in less-severe services is estimated to be over one million dollars. HYDROCRACKING CATALYSTS

Hydrocrackers normally include hydrotreating and hydrocracking sections. The hydrotreating catalyst is used in the pretreater to remove nitrogen, metals and sulfur contaminants from the feedstock. Nitrogen compounds are temporary poisons to downstream hydrocracking catalysts and can inhibit hydrocracking reactions. Thus, nitrogen compounds should be removed in the hydrocracker pretreater. Hydrotreating catalysts used in the hydrocracker pretreater are NiMo or NiW supported on gamma alumina. Hydrocracking catalysts are typically NiMo and NiW supported on amorphous silica-alumina or zeolite. Both the pretreating and hydrocracking catalysts must be sulfided to be activated. During this process, the pretreating and hydrocracking catalysts can be irreversibly deactivated by coke formation, metal poisoning, reduction of active metals and sintering. Nevertheless, the spent pretreating and hydrocracking catalysts can be regenerated to recover activity and reused. To be regenerable, the metal contaminants on the regenerated catalysts should not exceed the critical levels as listed in Table 1.

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SPECIALREPORT

CLEAN FUELS

als agglomeration that will result in reduced hydrogenation of the products. This phenomenon should be considered in units where the product quality is a key performance parameter, e.g., units producing Jet A1. Most regenerations are conducted ex-situ. According to a survey, both NiMo and NiW hydrocracking catalysts can be regenerated. Regen hydrocracker cat. Table 5

summarizes the effects of properly regenerated hydrocracking catalysts. Economic justification. Whether to use the regenerated or fresh hydrocracking catalyst is an economic decision. The main factors to consider include: • Cost of regeneration and makeup catalyst • Cost of fresh catalyst • Impact of a shorter cycle on maintenance and production cost • Cost of reduced product yield and quality due to regeneration • Benefits of new hydrocracking catalyst due to a longer cycle life and higher yield. The cost or benefit for each factor is refinery specific and should be considered in the economic evaluation accordingly.

TABLE 6. Product yields and volumes comparison1 Fresh catalyst, vol%

First-time regenerated catalyst, vol%

TABLE 7. Effect of each regeneration on properties and performance of semi-regen naphtha reforming catalyst

C1, wt%

0.61

2.76

Factor

C2, wt%

0.66

2.91

Surface area

% loss < 5%

LPG

Activity

< 1°C

Cycle length

negligible

Crush strength

negligible

Kero

Reformate and H2 yield1

negligible

LDO

HDO

Bleed

Based on a hydrocracker feed rate of 29,500 bpd.

Hydrocracking example. A refinery

used regenerated hydrocracking catalyst in its hydrocracker. At the same conversion level of 99%, the catalyst cycle life was fixed at 19 months by adjusting the deasphalted oil (DAO) feedrate. The DAO feedrate was reduced by approximately 700 bpd for the first regenerated catalyst cycle, compared to the fresh catalyst cycle. The typical product yields for the fresh catalyst cycle and first regenerated catalyst cycle are shown in Table 6.

The change of reformate and H2 yield is negligible as long as SA is above the minimum required.

An economic evaluation was done to compare the economics of the fresh hydrocracking catalyst cycle vs. firsttime regenerated hydrocracking catalyst cycle. The economic factors considered included cost of fresh catalyst, regeneration, makeup catalyst, along with the product value and DAO processing cost. The economic evaluation indicates that the fresh hydrocracking catalyst outperformed the first-time regenerated hydrocracking catalyst with a savings of millions over a cycle life of 19 months. In this case, the fresh hydrocracking catalyst should be used instead of the regenerated catalysts. NAPHTHA REFORMING CATALYSTS

The semi-regen naphtha reforming catalysts are chlorided platinum and rhenium (Pt and Re) or chlorided tri-metal (containing Pt and Re) supported on gamma alumina. During the reforming process, these catalysts can be irreversibly deactivated by coke formation, active metal agglomeration, support sintering or metal poisoning. The semi-regen naphtha reforming catalyst can be fully regenerated to recover its activity. It is a common practice that semi-regen naphtha reforming catalyst is regenerated in-situ and can be reused in several cycles. The regeneration is done to remove coke, re-disperse Pt and Re, and establish the correct chloride balance to restore catalyst activity and stability. At present, there is no ex-situ regeneration facility to regenerate the spent semi-regen naphtha reforming catalysts. Industrial practices. The number of regenerations that the refining industry conducts on the semi-regen naphtha reforming catalysts varies between 3 to 28 times with an average of 10–12 times. It is recommended that refiners dump, screen, and reload catalyst every third regeneration. All regenerations are conducted in-situ and all semi-regen naphtha reforming catalysts, Select 165 at www.HydrocarbonProcessing.com/RS 66

CLEAN FUELS

TABLE 8. Comparison of typical regeneration frequency

End of catalyst cycle

Do the regenerated catalyst SA, contaminant metal level and crush strength meet the requirements?

Yes

Catalyst type

Refining industry

A Middle East oil company

Hydrotreating

One to two times

Zero

Pretreating for the hydrocracker

Zero to one time in lesssevere services

Zero

Zero to one time

Up to 3 times

10–12 times

Up to 10 times

Hydrocracking Semi-regen reforming

Reclaim Pt of semi-regenerated reforming catalyst or sell the spent hydroprocessing catalyst for metal reclamation. Use fresh catalyst.

Does economic evaluation favor regenerated catalyst?

Most refiners replace their semiregen catalyst when the catalyst is unable Yes to provide the perforRegenerate the mance of the fresh catspent catalyst for reuse. alyst, mainly because the SA of the catalyst is below the miniFIG. 1 Decision-making chart for using regenerated catalysts. mum requirement to maintain good catalyst activity; or a new i.e., bi-metal and tri-metal, can be success- catalyst has been developed that can provide fully regenerated. It is a rule of thumb that more value than the existing catalyst. the regenerated semi-regen naphtha reforming catalyst should have the minimum sur- Industry vs. company practices. face area (SA) to be re-utilized economi- Table 8 summarizes utilization practices cally. The minimum SA required for good of regenerated refining process catalysts metal dispersion and chloride retention is between the refining industry and a major between 120 m2/g to 150 m2/g, depending Middle East oil company. on the vendor. The company is in agreement with the refining industry worldwide in using Effect of regenerated catalyst. If regenerated semi-regen naphtha reforming the catalyst is properly regenerated, the catalyst. However, the company doesn’t use effects of each regeneration are shown in regenerated hydrotreating catalyst as most Table 7. of the refining industry does. In the area Please note that SA loss after each regen- of hydrocracking catalyst, the company eration is not linear. The SA loss is high regenerates it up to three times, compared initially and it levels out as the number of with the zero to one time that the refining regenerations increases. The catalyst activ- industry does normally. ity can be sustained as long as the SA of the catalyst is above the minimum required to Decision making. The decision-makensure a good dispersion of Pt and Re. ing chart for using regenerated catalysts is shown in Fig. 1. At the end of the catalyst Economic justification. Whether to cycle, if the regenerated catalyst properuse the regenerated or the fresh semi-regen ties, i.e. surface area, contaminant metal naphtha reforming catalyst is an economic level and crush strength, can meet the decision. The main factors to consider requirements set by the catalyst vendor, we include: proceed with the economic evaluation to • Cost of fresh catalyst determine whether the fresh or the regen• Precious metal lease cost during pre- erated catalyst should be used. If not, use cious metal reclamation fresh catalyst, and reclaim the Pt of the • Production loss due to regeneration semi-regen naphtha reforming catalyst or • Impact of a shorter cycle on mainte- sell the spent hydroprocessing catalyst to a nance and production cost metal reclaimer. • Cost of reduced reformate and hydroIf the economic evaluation favors the gen yields regenerated catalyst, regenerate the spent • Benefits of new semi-regen naphtha catalyst for re-use in the future. If not, use reforming catalyst due to a longer cycle life the fresh catalyst, and reclaim the Pt of the and higher products yield. semi-regen naphtha reforming catalyst or No

SPECIALREPORT

sell the spent hydroprocessing catalyst to a metal reclaimer. Recommendations. The comparison of practices between the refining industry and a major Middle East oil company indicates that the company could use regenerated hydrotreating, and possibly hydrocracker pretreating catalysts, in less-severe applications, and regenerate the hydrocracking catalyst up to one time. The optimal number of regenerations of each refining process catalyst should be determined by a comprehensive economic evaluation that considers specific situations of each refinery. Since the base metal (Co, Ni, W and Mo) prices have been rising in recent years, the metal reclaimers are interested in purchasing spent hydrotreating and hydrocracking catalysts for metal reclamation. To protect the environment, it is recommended that the company’s refineries should make the sale of the spent hydrotreating and hydrocracking catalysts to metal reclaimers for metal reclamation or to catalyst vendors for regeneration, over disposal in a landfill if these catalyst are not suitable to be reused in the company’s processing units. HP

Co Mo Ni Pt W Re

NOMENCLATURE Cobalt Molybdenum Nickel Platinum Tungsten Rhenium

Dr. Gene J. Yeh is a registered professional engineer in the state of Louisiana. He holds a BS degree in chemical engineering and an MS degree and a Ph.D. in chemical and fuels engineering. Dr. Yeh has more than 22 years of experience in oil refining, catalyst manufacturing and R&D environment. He is currently working as an engineering specialist in the process & control systems department of Saudi Aramco in Dhahran, Saudi Arabia. His support areas include catalyst and adsorbent selection, hydrogen plant, hydroprocessing and naphtha reforming. HYDROCARBON PROCESSING FEBRUARY 2010

I 67

Select 83 at www.HydrocarbonProcessing.com/RS

HEAT TRANSFER

Calculating turbulent diffusion jet flames Following these steps may prevent potential damaging hot spots J. P. MEAGHER, Praxair, Inc., Tonawanda, New York

o estimate the impact of new operating conditions on processes and equipment under supervision, most engineers maintain a file of shortcut equations and spreadsheets to model equipment including heat exchangers, compressors and reactors. Unfortunately, one area that is often handled superficially with a rule-of-thumb or simple ratio is the flame profile estimation within fired heaters. The problem with these methods is that accuracy decreases with the separation from the base-case conditions. This can lead to missing important heat transfer effects and material limits within the fired heater. This usually results in destructive hot spots on the heater tubes or refractory walls. To address this inaccuracy, a method will be described for estimating the profile of a common type of flame: a momentum-dominated, turbulent diffusion jet flame in a co-flowing oxidant stream. Based on turbulent jet theory, the method allows the flame profile to be estimated for a variety of operating conditions. While this method does not reach the accuracy level obtained by field testing or computational fluid dynamics (CFD), it does provide quick and reasonable estimates that can be used in the prediction of the fired heater performance. As with any other engineering tool, accurate estimation of equipment performance can ultimately lead to increased operating profits and reduced maintenance costs. Flame identification. To ensure that the method described

is used for the proper flame type, the first step is to identify the characteristics of a momentum-dominated turbulent jet diffusion flame. To start, consider a candle flame in a draft-free room as shown in Fig. 1A. This flame has smooth edges and an almost steady appearance as the fuel and oxidant slowly enter the combustion zone. This is a laminar flame. In contrast, the flame from a handyman’s propane torch is shown in Fig. 1B. The fuel enters the combustion zone at a high velocity and many small-scale perturbations are generated. This is a turbulent flame. Now consider how the oxidant (air) reaches the fuel. For the candle, the oxidant and fuel diffuse toward each other through the flame boundary to support combustion.1 This makes it a diffusion flame. The fact that the fuel source is initially a solid is not important. In contrast, consider the handyman’s propane torch. If one examines this burner, there are small holes in the outer housing base from which no fuel normally flows. These holes allow the aspiration of small amounts of air into the fuel stream before it reaches the combustion zone. This creates a pre-mixed flame, even though the remainder of the fuel and air diffuse toward each other through the flame boundary.

FIG. 1

Common flame types.

If the aspiration holes were completely covered, the handyman’s torch would produce a diffusion flame because the oxidant and fuel would reach each other exclusively by diffusion through the flame envelope. Normally a handyman’s torch has a large fuel exit area (relative to the fuel nozzle) so that the pre-mixture of fuel and air can flow with little pressure drop. This large exit area results in a laminar flame when the aspiration holes are blocked. If the jet exit diameter were reduced at the same time as the aspiration holes were blocked, a turbulent diffusion flame would be produced, as shown in Fig. 1C. When considering the flames previously described, there is one more observation worth noting. If the candle is tilted, the flame would turn upward immediately. This is an example of a buoyant flame. If the handyman’s torch is tilted, only the flame tip would turn upward. This stiff flame is an example of a momentumdominated flame. This is the regime of interest. If there is a doubt regarding the momentum, refer to the referenced literature which describes how to confirm operation in the momentum-dominated region using the flame Froude number.1 Flame profile calculation. Using the turbulent diffusion jet flame theory in the cited literature, the steps for estimating the time-averaged flame profile in a co-flowing oxidant will be outlined and followed by several examples. In the description, the nomenclature from the original literature has been retained for the sake of continuity and the steps have been intentionally kept short to facilitate their use in a spreadsheet program. In addition, a HYDROCARBON PROCESSING FEBRUARY 2010

I 69

HEAT TRANSFER TABLE 1. Assumptions for turbulent diffusion flame analysis1

TABLE 3. Experimental conditions9

Fuel jet is turbulent at the source Fuel jet source is a circular opening

Composition, volume fraction N2/CH4/O2

Fuel jet flow is symmetrical about the central axis

Velocity, m/s

Fuel jet and oxidant flows are steady with initially uniform and parallel velocity profiles

Source diameter, mm

Fuel jet

Reynolds number

Co-flowing stream

0.7/0.3/0

0/0/1

0.3

150

18,600

Not reported

There is no global swirl No premixing of fuel and oxidant prior to combustion zone Fuel jet and resulting flame are dominated by momentum

= x+

The fuel jet is a free jet (does not interact with other fuel jets, flames, walls or other structures) Sufficient oxidant is available to prevent kinetic limitation Reaction zone resembles a thin sheet Turbulent transport is much greater than molecular transport of momentum, thermal energy and reacting species

TABLE 2. Experimental conditions8 Fuel jet Composition, volume fraction N2 + Ar /natural gas/O2

Co-flowing stream

0.6/0.4/0

0.79/0/0.21

Velocity, m/s

134

Source, mm

4.57 diameter

300 x 300 square

37,500

Not reported

Reynolds number

number of simplifications are required to make the entire analysis manageable. The primary assumptions are noted in Table 1. Step 1. Assuming that the fuel source is oxygen-free and the oxidant source is fuel-free, calculate the stoichiometric mixture fraction, XS , using the molar fraction of oxygen in the oxidant, x oxidant _ O , the molar fraction of each fuel species, xfuel_i , and the 2 positive multiple for the stoichiometric oxygen amount reacted per mole of fuel component i, â?ˇi , using Eq. 1.2 x oxidant _ O 2 Xs = (1) n x oxidant _ O + i=1 x fuel _ i i 2

(

)

Step 2. Determine an effective density, â?łâ&#x2C6;&#x17E;eff , for the surrounding oxidant to incorporate the reaction heat release effect.3 Use the stoichiometric mixture fraction, XS , just calculated, the oxidant source density, â?łâ&#x2C6;&#x17E;; the oxidant source temperature, Tâ&#x2C6;&#x17E; , stoichiometric reaction temperature, TS , for the fluids under consideration and the fuel source temperature, TE , with Eq. 2.

eff =

T TS TE T + E (1 X S )

(2)

Step 3. Calculate the modified axial distance from the fuel source, â?°. This includes the actual downstream distance from the nozzle opening plus an offset to consider the entire jet flow as its far-field equivalent.4 Use the actual distance downstream of the jet source, x, fuel jet mass flow, mE ; the jet exit velocity, UE ; the drag of the co-flowing oxidant stream on the fuel jet nozzle, D, and the effective density with Eq. 3. The drag may be calculated, as described in reference 5, but for a first approximation with typical low velocities of the co-flowing oxidant, the drag may be neglected. 70

I FEBRUARY 2010 HYDROCARBON PROCESSING

7.25mE ( eff(mEU E

D))

(3)

1/2

Step 4. Find the intermediate value called the extended momentum radius, â?Ş+, from reference six, using previously defined parameters and the co-flowing oxidant source velocity, Uâ&#x2C6;&#x17E; , with Eq. 4. As before, the drag may be neglected for low co-flowing oxidant velocities. m U D + = E Eeff 2 U

1/2

(4)

Step 5. Calculate the blending function that incorporates the full range of flow behavior from the jet-like limit, where the fuel stream velocity dominates, to the wake-like limit, where the oxidant velocity dominates using Eq. 5.2,6 g = 2,915 + + + +

1/ 3

44.2 + 3.31 +

(5)

Step 6. Calculate the stoichiometric mixture fraction, â?¨S , using parameters previously defined with the molecular weights of the oxidant and fuel, MWoxidant and MWfuel respectively, from Eq. 6.2 MWoxidant s = 1+ MW xoxidant _ O fuel 2

(

)(

)

(

1 i=1 x fuel _ i i (6) n

)

Step 7. Calculate the radial distance from the centerline, r, where the mixture reaches the stoichiometric value using Eq. 7.4,6 This radial distance is the location of the flame edge for axial distances downstream of the jet source that are less than or equal to the flame length.4,6 This radial value approaches zero as the 1/2 flame length is reached.

(m U D)

S +

E E 2.1

ln r( ) = (7) 2

/3 1.8m U E

44.2 + 3.31

+ As a check, the method just outlined was used with the conditions from two turbulent diffusion jet flame studies in co-flowing oxidants where the jets were free or nearly free as determined by the Craya-Curtet number.7 In the first study by MuĂąiz and Mungal, a fuel mixture of natural gas and nitrogen was combusted in a co-flowing air stream.8 Experimental conditions for the high Reynolds number flame are noted in Table 2. A flame length to jet exit diameter ratio of 152 was reported. The method described here yielded a value of 168 with a flame width to jet exit diameter

PROCESS INSIGHT Comparing Physical Solvents for Acid Gas Removal Physical solvents such as DEPG, NMP, Methanol, and Propylene Carbonate are often used to treat sour gas. These physical solvents differ from chemical solvents such as ethanolamines and hot potassium carbonate in a number of ways. The regeneration of chemical solvents is achieved by the application of heat whereas physical solvents can often be stripped of impurities by simply reducing the pressure. Physical solvents tend to be favored over chemical solvents when the concentration of acid gases or other impurities is very high and the operating pressure is high. Unlike chemical solvents, physical solvents are non-corrosive, requiring only carbon steel construction. A physical solvent’s capacity for absorbing acid gases increases signiﬁcantly as the temperature decreases, resulting in reduced circulation rate and associated operating costs.

PC (Propylene Carbonate) The Fluor Solvent process uses JEFFSOL® PC and is by Fluor Daniel, Inc. The light hydrocarbons in natural gas and hydrogen in synthesis gas are less soluble in PC than in the other solvents. PC cannot be used for selective H2S treating because it is unstable at the high temperature required to completely strip H2S from the rich solvent. The FLUOR Solvent process is generally limited to treating feed gases containing less than 20 ppmv; however, improved stripping with medium pressure ﬂash gas in a vacuum stripper allows treatment to 4 ppmv for gases containing up to 200 ppmv H2S. The operating temperature for PC is limited to a minimum of 0°F (-18°C) and a maximum of 149°F (65°C).

Gas Solubilities in Physical Solvents All of these physical solvents are more selective for acid gas than for the main constituent of the gas. Relative solubilities of some selected gases in solvents relative to carbon dioxide are presented in the following table. The solubility of hydrocarbons in physical solvents increases with the molecular weight of the hydrocarbon. Since heavy hydrocarbons tend to accumulate in the solvent, physical solvent processes are generally not economical for the treatment of hydrocarbon streams that contain a substantial amount of pentane-plus unless a stripping column with a reboiler is used.

Typical Physical Solvent Process

Gas Component

DEPG at 25°C

PC at 25°C

NMP at 25°C

MeOH at -25°C

DEPG (Dimethyl Ether of Polyethylene Glycol)

0.013

0.0078

0.0064

0.0054

DEPG is a mixture of dimethyl ethers of polyethylene glycol. Solvents containing DEPG are marketed by several companies including Coastal Chemical Company (as Coastal AGR®), Dow (Selexol™), and UOP (Selexol). DEPG can be used for selective H2S removal and can be conﬁgured to yield both a rich H2S feed to the Claus unit as well as bulk CO2 removal. DEPG is suitable for operation at temperatures up to 347°F (175°C). The minimum operating temperature is usually 0°F (-18°C).

Methane

0.066

0.038

0.072

0.051

Ethane

0.42

0.17

0.38

0.42

CO2

1.0

Propane

1.01

0.51

1.07

2.35

n-Butane

2.37

1.75

3.48

COS

2.30

1.88

2.72

3.92

MeOH (Methanol)

H 2S

8.82

3.29

10.2

7.06

The most common Methanol processes for acid gas removal are the Rectisol process (by Lurgi AG) and Ifpexol® process (by Prosernat). The main application for the Rectisol process is puriﬁcation of synthesis gases derived from the gasiﬁcation of heavy oil and coal rather than natural gas treating applications. The two-stage Ifpexol process can be used for natural gas applications. Methanol has a relatively high vapor pressure at normal process conditions, so deep refrigeration or special recovery methods are required to prevent high solvent losses. The process usually operates between -40°F and -80°F (-40°C and -62°C).

n-Hexane

11.0

13.5

42.7

Methyl Mercaptan

22.4

27.2

34.0

NMP (N-Methyl-2-Pyrrolidone) The Purisol Process uses NMP® and is marketed by Lurgi AG. The flow schemes used for this solvent are similar to those for DEPG. The process can be operated either at ambient temperature or with refrigeration down to about 5°F (-15°C). The Purisol process is particularly well suited to the purification of high-pressure, high CO2 synthesis gas for gas turbine integrated gasification combined cycle (IGCC) systems because of the high selectivity for H2S.

Choosing the Best Alternative A detailed analysis must be performed to determine the most economical choice of solvent based on the product requirements. Feed gas composition, minor components present, and limitations of the individual physical solvent processes are all important factors in the selection process. Engineers can easily investigate the available alternatives using a veriﬁed process simulator such as ProMax® which has been veriﬁed with plant operating data. For additional information about this topic, view the technical article “A Comparison of Physical Solvents for Acid Gas Removal” at http://www.bre.com/tabid/147/Default.aspx. For more information about ProMax, contact Bryan Research & Engineering or visit www.bre.com.

Bryan Research & Engineering, Inc. P.O. Box 4747 • Bryan, Texas USA • 77805 979-776-5220 • www.bre.com • sales@bre.com Select 113 at www.HydrocarbonProcessing.com/RS

HEAT TRANSFER TABLE 4. Summary of operating parameters used in Fig. 2

4 Methane and exhaust Methane and air Hydrogen and air

Axial distance, m

Case

Fuel

Methane

Hydrogen

Methane

Flow, kg/h

2.1

Exit velocity, m/s

103

341

103

Stagnation temperature, °C

Oxidant

Air

Exhaust

Velocity, m/s

0.7

0.6

Stagnation temperature, °C

100

Oxygen, % by volume

the oxidant with fuel, typical of staged burners in low-emissions applications, has a tendency to increase the flame length in the 0 same manner. -2 -1 0 1 2 A step-by-step approach for calculating the envelope of a Radial distance, m momentum-dominated turbulent diffusion jet flame in a coflowing oxidant was discussed. While it was shown that reasonFIG. 2 Turbulent diffusion flame profiles for different fuels and oxidants. able results can be readily obtained, some practical limitations should be remembered for actual equipment. The first is that ratio of 22. Since actual mixture fraction variations result from the the correlations were developed from observations of jet flames turbulent nature of the flows involved, flame length fluctuations on having flickering edges so the existence of some variation must the order of the flame width should be anticipated.4 The conclube accommodated. This includes the fact that, no matter the presion is that this estimate is reasonable. In another study for a high dicted flame envelope, there are extremely hot gases in the vicinity. Reynolds number flame, a fuel jet of methane plus nitrogen was The second is that the flame blowout is not identified with this combusted in a co-flowing stream of pure method and should be checked separately. oxygen.9 These conditions are summarized ■ Since actual mixture Finally, any flame lift-off from the nozzle in Table 3. Results from the study indiprovides some premixing of oxidant into cated a flame length to jet exit diameter fraction variations result from the fuel jet base, which will shorten an ratio of 48. Using the method described, actual flame from what was predicted. a flame length to jet exit diameter ratio of the turbulent nature of the 60 was estimated with a flame width to jet A practical method was flows involved, flame length Conclusion. exit diameter ratio of 11. Considering the outlined to estimate the profile of a typical flame appearance with indistinct fluctuations on the order of momentum-dominated turbulent diffuand flickering edges, these correlations sion jet flame in a co-flowing oxidant. The provide acceptable results. intent is to enable the use of simplified the flame width should be As another illustration of the insight jet theory for the investigation of a wide offered by the approach outlined here, anticipated. range of combustion parameters rather flame profile estimates for three differthan depending on a simple ratio or ruleent fuel/oxidant cases are shown in Fig. 2. These are methane/ of-thumb. Realistic burner performance estimates are a key part air, hydrogen/air and methane/exhaust gas. On the vertical scale of confirming reliable heater operation at greater capacity or of Fig. 2, the axial distance from the source is shown, and on the preventing potential damaging hot spots when new conditions horizontal scale, the local flame width is indicated. These scales are used. Either can help increase the profitability of the process are of a similar magnitude to offer some perspective on the flame equipment under an engineer’s supervision. HP proportions. Operating parameters for these cases are summaACKNOWLEDGMENTS rized in Table 4. Each assumes the same net heat release rate, fuel The author thanks Dr. Werner Dahm of the University of Michigan and and oxidant source diameters and neglects the drag force of the Dr. Lee Rosen of Praxair, Inc. for their expertise and assistance in preparing this co-flowing oxidant. With the methane/air flame as the basis for article. comparison, note the reduction in flame length when the fuel is LITERATURE CITED switched to hydrogen. For jet flames that are not limited by the 1 Turns, S. R., An Introduction to Combustion: Concepts and Applications, amount of surrounding oxidant, this length reduction is due to Second Edition, WCB/McGraw-Hill, Boston, 2000. the greater entrainment rate of the lower density fuel.10 Using 2 Dahm, W. J. A., Jet Physics and Theory for Combustion Applications Involving the methane in air flame profile shown in Fig. 2 as the basis for Free or Confined Jets, University of Michigan, Ann Arbor, 2006. 3 Tacina, K. M. and W. J. A. Dahm, Journal of Fluid Mechanics, Vol. 415, pp. comparison again, note that when the oxidant is changed from 23–44, 2000. air to a warm, oxygen-reduced exhaust gas, the flame lengthens. 4 Diez, F. J. and W. J. A. Dahm, Journal of Fluid Mechanics, Vol. 575, pp. It is caused by the higher co-flowing velocity, which reduces mix221–255, 2007. ing, and the lower concentration of oxygen, requiring a greater 5 Lindeburg, M. R., Mechanical Engineering Reference Manual for the PE Exam, distance to reach the stoichiometric value. Delayed mixing of 2001. 72

I FEBRUARY 2010 HYDROCARBON PROCESSING

HEAT TRANSFER 6

Nagel, Z. W. and W. J. A. Dahm, Experimental Study of Heat Release Effects in Exothermically Reacting Turbulent Shear Flows, Laboratory for Turbulence & Combustion, Department of Aerospace Engineering, University of Michigan, Ann Arbor, Michigan, 2007. 7 Guruz, A. G., H. K. Guruz, S. Osuwan and F. R. Steward, Combustion Science and Technology, Vol. 9, pp. 103–110, 1974. 8 Muñiz, L. and M. G. Mungal, Combustion and Flame, Vol. 126, pp. 1402– 1420, 2001. 9 Donbar, J. M., J. F. Driscoll and C. D. Carter, Combustion and Flame, Vol. 125, 1239–1257, 2001. 10 Ricou, F. P. and D. B. Spalding, Journal of Fluid Mechanics II, pp. 21–32, 1961.

XS S + E ∞

∞eff i

Stoichiometric mixture molar fraction Jet half-width at 50% of the centerline velocity, length Stoichiometric mixture fraction of conserved scalar Extended momentum radius, length Modified axial distance from jet source, length Jet exit density, mass/length3 Density of co-flowing oxidant stream at source, mass/length3 Effective density of co-flowing oxidant considering heat release, mass/length3 Stoichiometric moles of oxygen reacted per mole of fuel component i

NOMENCLATURE The following parameters were used in this article. Any consistent set of units may be applied. dE D g L mE MWfuel MWoxidant r TE TS T∞ UE U∞ x xfuel_i

x oxidant _ O

Jet exit diameter, length Drag of co-flowing stream on outside of fuel nozzle, force Blending function Flame length, length Jet exit mass flow, mass/time Molecular weight of the fuel stream, mass/mole Molecular weight of the oxidant stream, mass/mole Radial distance from jet axis, length Fuel jet exit temperature, absolute scale Stoichiometric reaction temperature, absolute scale Co-flowing oxidant source temperature, absolute scale Jet exit velocity, length/time Co-flowing velocity of the oxidant fluid, length/time Axial distance from jet source, length Molar fraction of each fuel species i Molar fraction of oxygen in the oxidant

James P. Meagher is a senior development associate at Praxair’s research and development department in Tonawanda, New York. He has 28 years of process and equipment experience ranging from cryogenic air separation and hydrogen liquefaction plants to high temperature combustion and synthesis gas generation systems. Mr. Meagher has a BS degree in mechanical engineering from the University of Notre Dame and recently completed his MS degree in chemical engineering at the State University of New York at Buffalo. He is also a registered Professional Engineer in the State of New York and an inventor on several US patents.

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Results

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INSTRUMENTATION

Delayed coker heater analysis using an artificial neural network The model was used to study the effects of the many variables that affect coke formation M. SHARMA, A. PONSELVA, Reliance Industries Limited, Jamnagar, India

oking inside tubes is a major issue in delayed coker units. Due to coke formation inside tubes, they cannot be continuously operated after a few months. Thus, they have to be shut down and the coke has to be removed. The coke formation rates in heaters is a function of many variables. This article studies the effects of these variables on coke formation rate inside heaters. For further analysis an artificial neural network (ANN) model was used. It was found that the model could map the situation satisfactorily. The model results are presented. Introduction. The delayed coker heaters

in this study are pigged every five months (on average). The coking rate is a function of various variables. These variables are controlled on a trial-and-error basis to maximize coker heater run lengths based on the valuable experience of operators and engineers. The delayed coker has been operated with a wide variety of feeds and operating conditions and the responses to these variations have been captured in the form of IP 21 data. This database is a storehouse of knowledge that can be efficiently used to help engineers understand various heater operation modes and the impact of feed quality on heater run lengths. This helps to optimize heater operation and maximize run lengths. The feed quality data were obtained from daily reports supplied by the refinery. The following article discusses a procedure to capture this information in the form of an ANN model. Then using this model the effects of different variables are analyzed.

function is determined largely by the connections between elements. We can train a neural network to perform a particular function by adjusting the values (weights) of the connections between elements.

■ Neural networks have

been trained to perform complex functions in various fields, including pattern recognition, identification, classification, speech, vision and control systems. Neural networks have been trained to perform complex functions in various fields, including pattern recognition, identification, classification, speech, vision and control systems. Neural networks can be trained to solve problems that are difficult for conventional computers or humans. It has been proven mathematically that a two-layered neural network can model any continuous function. The average skin temperature was chosen as the output. The input variables are the same as that used for input variables.

ANN model. Neural networks are com-

posed of simple elements operating in parallel. These elements are inspired by biological nervous systems. As in nature, the network

FIG. 1

The model chosen had the following properties: • No. of layers = 2 • No. of neurons in the first layer = 3 • Transfer function of the first layer = “logsig” (sigmoid) • No. of neurons in the second layer = 1 • Transfer function of the second layer = “purelin” (linear) • Input range = 0 to 1 for all variables • Training function = “trainlm” • Adaptation learning function = “learngdm” • Performance function = mean-square error (MSE ) • Network type = feedforward back propagation. The ANN model is shown in Fig. 1. This model was arrived at after trials with different models. It should be noted that apart from accuracy in prediction, to prevent over-fitting the number of neurons was kept low. Model training, validation and testing. To make the model the collected

data set is divided into three parts: training, validation and testing. Training data sets: These are presented to the network during training, and the

IW{1.1}

LW{2.1}

b{1}

b{2} 3

An ANN model is composed of simple elements operating in parallel.

HYDROCARBON PROCESSING FEBRUARY 2010

I 75

INSTRUMENTATION 1.0

1.2 1.0 0.8 0.2 0.4 0.2 0.0

0.7 0.6 0.5 0.4 0.3 0.2

0.8 0.7 0.6 0.5 0.4 0.3 0.2

0.1

0.0

577

576.5 576.0 575.5 575.0 574.5 574.0 573.5 573.0 572.5

575.5 575.0 574.5 574.0 573.5 573.0 572.5 572.0

576 575 574 573 572 1

Skin temperatures (°C) vs. COT.

FIG. 4

network is adjusted according to its error. In other words, using an iterative procedure, the ANN model adjusts its weights and biases according to the data points. Validation data sets: These are used to measure network generalization, and to halt training when generalization stops improving. Testing data sets: These have no effect on training and provide an independent measure of network performance during and after training. Using all the data sets we can be sure that the model weights and biases are appropriate. For the present case the data were divided in the ratios: • 70% (114 data points) were used for training • 20% (33 data points) were used for validation • 10% (16 data points) were used for testing MSE and in prediction was less than 10–3 in all three data sets. To further evaluate the goodness for fit, the model was used to predict the output for all data points. Then a plot was drawn between the predicted and actual values. It was found that in all three data sets the data points lie on the y = x line. 76

0.8

Data points Best linear Fn A=T

0.9

Graphs of b/w actual values vs. ANN predictions for the training, validation and testing sets.

FIG. 2

FIG. 3

1.0 Data points Best linear Fn A=T

0.9

Output A, linear Fn A=(I)T+(0.006)

Data points Best linear Fn A=T

Output A, linear Fn A=(I)T+(0.034)

Output A, linear Fn A=(I)T+(0.014)

1.4

I FEBRUARY 2010 HYDROCARBON PROCESSING

Skin temperatures (°C) vs. VR flowrate.

575.0 574.8 574.6 574.5 574.2 574.0 1

FIG. 5

Skin temperatures (°C) vs. VR flow fluctuations.

FIG. 6

576.5 576.0 575.5 575.0 574.5 574.0 573.5 573.0 572.5

Skin temperatures (°C) vs. CIT.

The plots for different data sets are as shown in Fig. 2. Data analysis using ANN. The ANN

model was fitted for heater 1. Then the effect of each variable was studied. To study the effect of a particular variable, all other variables were maintained at mid values, i.e., (maximum+minimum)/2. Figs. 3 and 4 correspond to pass one. In pass three the boiler feed water (BFW) flowrate was varied in a very high range compared to the others hence, it has been studied separately. The x-axis shows the number of increments in the corresponding variable. The range in which the vari-

1 FIG. 7

Skin temperatures (°C) vs. BFW flowrate.

able has varied was divided by 10. One increment means the value is increased by 10% of the range. The value in y-axis shows the average skin temperature (°C) that would have been attained on the 82nd day if the value on the x-axis was maintained on the 82nd day. Effect of COT and VR flowrate. As suspected, the effect of these variables was most high and they had a positive influence on average skin temperature. Hence, these variables are not studied in detail. Effect of coil inlet temperature (CIT). As can be seen in Fig. 5, for all passes CIT is negatively influencing. Thus, by increasing CIT to the maximum of the present range

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INSTRUMENTATION the skin temperature on a particular day would have been approximately 3°C lower. Hence, this a very influential parameter to increase heater run length. Effect of VR flowrate fluctuations. As evident from Fig. 6, flowrate fluctuations also have an appreciable effect. Thus, decreasing VR flowrate fluctuations can be considered as a feasible way to increase run length. It is interesting to note that this graph has a curvature around the fourth point. In other words, the effect of an increase in VR flowrate fluctuations is more significantly felt when the fluctuation is less. When fluctuations become higher, the effect is comparatively less. Effect of BFW flowrate. It should be noted that in all the passes BFW flowrate

has been maintained between 300 to 350 kg/hr. It is interesting to note that in this range BFW injection doesn’t play a positive role in skin temperature reduction (Fig. 7). This phenomenon can be explained if we look into the heat transfer mechanics in the tube. BFW injection has two effects, on one hand it decreases residence time (by increasing velocity) but on the other hand it takes away appreciable heat from the VR flowing in the tube. The balance between these two effects decides the overall effect of BFW injection. Effect of feed quality. It’s quite interesting to note that CCR plays a slightly negative 577 576 575 574

575.9

573

575.4

572 571

574.9

574.4 FIG. 10

573.9

Temperature (°C) vs. cross-over temperature.

role in skin temperature rise (Fig. 8). Thus, in the present range of study, i.e., between 23 and 28, increasing CCR decreases the rate of skin temperature rise. This can be explained by taking into account resins. With an increase in CCR the resin content in the VR is increasing. It is well known that resins are required to keep asphaltenes stable. Thus, with an increase in CCR, the coking tendency of VR is reduced and hence, skin temperature rise is less. But asphaltenes play an opposite role than that of CCR (Fig. 9). With an increase in asphaltenes, there will be a consistent rise in skin temperature. This trend is consistent with present theories of coke formation. Effect of cross-over temperature. As shown in Fig. 10, increasing cross-over temperature decreases skin temperature. Thus, if it is possible to increase heat transfer into the convection section then run length can be increased. Comparing effects of different variables. To study the degree of influence of each variable, the range in which skin temperature has changed is a good indicator. Fig. 11 shows the range for each variable.

573.4 1

9 5

Temperature (°C) vs. CCR of VR.

FIG. 8

4 3 2

576.0

575.5

0 -1

575.0

-2

574.5

-3

574.0

-4

573.5

COT 1

9 FIG. 11

FIG. 9

Temperature (°C) vs. asphaltene content.

FIG. 12

Contour plot of skin temperature vs. COT and BFW flowrate.

I FEBRUARY 2010 HYDROCARBON PROCESSING

BFW ﬂow

CIT

VR ﬂow

VR ﬂuctuation

API

CCR

Asphaltene Cross overt

Comparison of degree of influence of variables.

FIG. 13

Contour plot of skin temperature vs. COT and asphaltene content.

INSTRUMENTATION A negative value indicates with an increase in the variable the skin temperature drops. The absolute value is an indicator of the degree of influence. As studied in the linear regression model, we find that COT and VR flowrate have the most influence, BFW flowrate has very little effect. It should be noted that heater operating variables have more effect than feed quality. In other words, CIT, cross-over temperature and VR flowrate fluctuations have more effect than that of feed quality. But it is clearly visible that the influence of feed quality is not as snall as suggested by the linear model. The histogram clearly shows that asphaltene content is quite appreciable and this is the most influencing variable among feed quality parameters. CCR and density also have appreciable effect, but less than heater operating variables. Variable interaction. As mentioned earlier, the operation range has been a little higher in pass three of heater one. Also, the linear models suggested the presence of interaction between variables like COT and BFW flowrate. Fig. 12 shows contour plot between BFW flowrate and COT. The

lighter regions indicate that skin temperature is low and in darker regions skin temperature is very high. In Fig. 12 we find in all regions COT increases skin temperature and BFW flowrate decreases skin temperature. Also, from the slope of regions we can say that COT is far more influential than BFW flowrate. It is interesting to note that BFW flowrate has no effect up to 390 kg/hr but after that there is a sudden change in curvature. This indicates beyond this range BFW flowrate is quite influencing and results in a drop of coke formation rate. The Fig. 13 contour plot between COT and asphaltene content shows that there is almost no interaction between these two variables in the present range of study. Both of them have their own independent effects. Both result in increased skin temperatures. The same trend is present in all combinations of the two variables. Conclusion. The above findings shows

that an ANN model can satisfactorily model the effect of variables influencing coke formation rate in coker heaters. The model shows that even though COT and

VR flowrate were varied in a very small range, they are the most influencing variables. Further, it was found that CIT and cross over temperature negatively influence skin temperature. Hence, increasing them will improve runlength. In feed quality, asphaltene content results in higher skin temperature, but CCR is found to decrease the rate of skin temperature rise. HP

Mukesh Sharma has an M.Eng. degree from the University of British Columbia, Canada. He previously worked with Reliance Industries Limited. Jamnagar, Gujarat as process technologist. Mr. Sharma has more than five years of experience in the process industries. He is an Associate member of Indian Institute of Chemical Engineers. His main interests are modeling and simulation, process design and project engineering.

Anto Ponselvan is a final-year postgraduate student at IIT Roorkee. His areas of interest are process modeling and simulation. He has published three research papers and participated in many conferences. He worked as an intern under Mr. Sharma at Reliance industries Limited, Jamnagar, Gujarat.

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INSTRUMENTATION

H max =

Qi Hvl g( min ) A

Level, %

where t = 0.404. In this case, the tightest available Q o max is 1.3606 Qi or equivalently min ^0.404. • For a given (Q'o max , Hmax ) set with h g , the PI controller for the tightest constraint control of Qo max can be obtained by designing it on vr, vrH. The tightest available Qo max is calculated by Qo max = Q i f ( vr) from Eq. 11 and the tightest available min becomes min = vr • For a given (Qo max , Hmax ) set, the PI controller that gives the tightest Q'o max can be designed based on ( min, vuH). The tightest available Q'o max is calculated by Qih ( min ) Q 'o max = Hvu . • For a given (Q'o max , Qo max ) set, the PI controller that gives the tightest Hmax can be obtained using ( min, vlH). The tightest available Hmax is obtained as

50 45 0.0

Illustrative examples. Consider a liquid level loop: A = 1

m2, Hspan = 2 m, Qov max = 4 m3/min, Q'o max = 5 m3/min2. The

0.5

1.0

1.5

2.0 2.5 Time, min

3.0

3.5

4.0

0.5

1.0

1.5

2.0 2.5 Time, min

3.0

3.5

4.0

0.5

1.0

1.5

2.0 2.5 Time, min

3.0

3.5

4.0

5 4 3 2.3 2 1 0 0.0 1.5 1.3 1.0 0.5 0.0 0.0

Example 1 Example 2 Example 3 Example 4

60 57.5 55

dQo /dt (m3/min2)

PI controller design for tightest constraint control.

For a level loop being operated near the constraint, it is necessary to control the loop on the tightest possible constraint, the socalled “tightest constraint control”: • For a given Hmax (or Q'o max ), the PI controller for the tightest constraint control with respect to Q'o max (or Hmax ) can be designed using either g H = or H = h h( t ) t g( )

initial steady-state level is assumed to be at 50%. The expected maximum Qi is 1 m3/min, and w is chosen as 0.8. Seven examples are investigated to illustrate the corresponding seven possible cases. The constraint sets and the resulting optimal PI parameters are listed in Table 2. Figs. 5a and 5b show the responses afforded by the resulting optimal PI controller. In the simulation, a step change in the inflow is made at t = 1. As seen

ΔQo (m3/min)

Part 1 of this article outlined the methodology. This part will discuss PI controller design for tightest constraint control and provide some examples.

FIG. 5A Liquid level loop system responses for examples 1, 2, 3 and 4. HYDROCARBON PROCESSING FEBRUARY 2010

I 81

INSTRUMENTATION TABLE 2. Constraint specifications and PI parameters for the illustrative examples

Level, %

65 Example 5 Example 6 Example 7

60 57.5 55

Example Case

dQo /dt (m3/min2)

45 0.0

0.5

1.0

1.5

2.0 2.5 Time, min

3.0

3.5

4.0

5 4 3 2 1 0 0.0

0.5

1.0

1.5

2.0 2.5 Time, min

3.0

3.5

1.50 ΔQo (m3/min)

0.50 0.00 0.0

0.5

1.0

1.5

2.0 2.5 Time, min

3.0

3.5

from the figures, the level loop is optimally controlled and strictly satisfies the given three constraints in every case. It should be noted that the PI parameters for examples 4, 6 and 7 also imply those for the tightest constraint control with respect to Qo max , Q'o max , and Hmax when the (Q'o max , Hmax ), (Qo max , Hmax ) and (Q'o max , Qo max ) set are assumed to be given as those of examples 4, 6 and 7 in Table 2, respectively. The tightest Qo max , Q'o max , and Hmax specifications are calculated as 1.2476 m3/min, 4.7746 m3/min2 and 0.2387 m, respectively. The responses of examples 4, 6 and 7 in Figs. 5a and 5b clearly meet these tightest constraints. HP End of series: Part 1, January 2010. ACKNOWLEDGMENT This research was supported by Yeungnam University research grants in 2008.

2.3

1.0

1.3

1.15,

0.46

5.0

0.15

1.3

2.3601,

0.6796

4.0

0.15

1.3

2.0,

0.3635

5.0

0.3

1.15

1.8766,

0.9111

6.0

0.15

1.15

2.3873,

0.7162

3.0

0.5

1.15

1.5,

1.1398

NOMENCLATURE Roman letters Drum cross-sectional area, m2 Level deviation from setpoint, m Level transmitter span, m Maximum allowable level deviation from setpoint, m Proportional gain (dimensional), m2/min Proportional gain (dimensionless) Inflow change, m3/min Outflow change, m3/min Maximum allowable change in outflow, m3/min Maximum achievable outflow through the outlet valve, m3/min Expected maximum step change in inflow, m3/min Rate of change in outflow, dQ'o(t)/dt, m3/min2 Maximum allowable rate of change in outflow, m3/min2 Maximum achievable rate of change in outflow, m3/min2 Time, min

I FEBRUARY 2010 HYDROCARBON PROCESSING

KC 0.6325,

Drum surge volume, A⌬Hspan, m3

␴i

Greek letters Damping factor Integral time constant, min Apparent holdup time, A /KL or ␶V /Kc , min Holdup time, A⌬Hspan / Qovmax , min Weighting factor of objective function, 0 < ␻ < 1 Objective function or performance measure for optimal control Lagrangian multiplier Slack variable

t † *h, *g, *f

Superscripts Tangent point Extreme point of the objective function Optimum point on the constraint

␻i

4.0

FIG. 5B Liquid level loop system responses for examples 5, 6 and 7.

A H ⌬Hspan Hmax KL Kc Qi Qo Qomax Qovmax ⌬Qi Q'o Q'omax Q'ovmax t

␨ ␶I ␶H ␶V ␻ ␾

1.15 1.00

Qo max 1.3

␶I 1.5811

4.0

Hmax 1.0

PI parameter

Specification

Q 'o max 4.0

H = h h( ), H =

g( )

= f ( ),

respectively vr, vu, vl

Optimum point on the vertex formed by

H =

g( ) = f ( ),

and

H = h h( ), H = g( ) and f H = h h( ) and f = f ( ), respectively

LITERATURE CITED Buckley, P. (1983), “Recent advances in averaging level control,” Productivity through Control Technology, 18–21, Houston. 2 Cheung, T. and Luyben, W. (1979), “Liquid-level control in single tanks and cascades of tanks with P-only and PI feedback controllers,” Ind. Eng. Chem. Fund., 1, 15–21. 3 Luyben, W. and Buckley, P.S. (1977), “A proportional-lag controller,” Instrum. Tech., 24, 65–68. 4 MacDonald, K., McAvoy, T. and Tits, A. (1986), “Optimal averaging level control,” AIChE J., 32, 75–86. 5 Marlin, T. E. (1995), Process Control: Designing processes and Control Systems for Dynamic Performance, McGraw-Hill Inc., New York, USA. 6 Rivera, D.E., Morari, M. and Skogestad, S. (1986), “Internal model control. 4. PID controller design,” Ind. Eng. Chem. Process Des. Dev., 25, 252–265. 7 Seki, H. and Ogawa, M. (1998), Japan Patent # 2811041. 8 St. Clair, D.W. (1993), Controller tuning and control loop performance:PID without math, 2nd ed., Straight-Line Control Co Inc., Newark, Delaware. 9 Wu, K., Yu, C. and Cheung, Y. (2001), “A two degree of freedom level control,” J. of Process Control, 11, 311–319. 10 Shin, J., Lee, J., Park, S., Koo, K. and Lee, M., “Analytical design of a proportional-integral controller for constrained optimal regulatory control of inventory loop,” Com. Eng. Prac., Vol. 16, pp. 1391–1397, 2008. 11 Lee, M. and Shin, J., “Constrained optimal control of liquid level loop using a conventional proportional-integral controller,” will appear in Chem. Eng. Comm., 2009. 12 Dennis, J.B. (1959), Mathematical Programs and Electron Networks, Jhon Wiley & Sons, New York. 1

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INSTRUMENTATION Appendix. Derivation of performance measure and inequality constraints given in Eqs. 11aâ&#x20AC;&#x201C;d

Derivation of the performance measure â?ž given by Eq. 11a10 For a step change of magnitude â&#x152;ŹQi in the inflow, H(t) and Q'o(t) are obtained from Eqs. 4 and 5, respectively, as follows: Qi e r1t e r2 t H (t ) = (A1) for r1 r2 A r1 r2 r 2 e r2 t r 2e r1t for r r 1 (A2) Q 'o (t ) = Qi 2 1 2 r r 1 2 where r 1 and r 2 are the roots of the characteristic equation H I s2 + I s +1= 0 . Therefore, the performance measure for optimal control is: 2 2 H (t ) Q ' (t ) o dt dt + (1 ) 0 Q ' 0 H span

ov max

2 Q 1 2 2 1 1 1 i = + + A H span r1 r2 r1 + r2 2 r1 r2

2 2 Q 1 r 3 r 3 2r 2r 2 i 2 1 + 1 2 (1 ) 2 r1 + r2 Q 'ovmax r1 r2 2 2 2 Q

i 3 2+ (1 ) Qi = 2 H Q ' A H span ov max 1 1 1+ 2 H 4 2

(A3)

Derivation of the constraint given by Eq. 11d The outflow response to a step change in the inflow is obtained from the inverse Laplace transform of Eq. 5 as follows:

Qo (t ) = Qi 1+ exp sin( t )

cos( t ) +

2 H 1 2

for < 1 where =

1 2 H

1 2

= Qi 1+ for = 1

exp

+1 2 H 2 H

cosh( t ) + sinh( t ) = Qi 1+ exp

2 2 H

1 for > 1 where =

1 2 H

2 1 (A4)

The peak of Qo(t) can be found from dQo(t)/dt = 0 in terms of â?¨: Qo peak = Qi f ( ) (A5) where f (â?¨) is:

I FEBRUARY 2010 HYDROCARBON PROCESSING

2 tan 1 x f ( ) = 1+ exp x 1 2 = 1+ exp( 2) for = 1 2 tanh 1 x = 1+ exp x for 0 < < 1 where x =

for > 1 where x =

2 1

(A6)

Therefore, the constraint in Eq. 11d can be expressed as: Qi f ( ) Q o max

(A7)

11b11

Derivation of the constraint given by Eq. The rate of change of the outflow, Q'o(t), is obtained by differentiating Eq. A4. The peak of Q'o(t) can be found from dQ'o(t) /dt in terms of â?¨ and â?śH : Qi Q 'o peak = h( ) (A8) H where h(â?¨) is: h( ) =

3 tan 1 x 1+ x 2 exp 2 x

for 0 < < 0.5 where x =

1 2

= 1 for 0.5

(A9)

Therefore, the constraint given by Eq.11b can be expressed as: Qi h( ) Q 'o max (A10) H Derivation of the constraint given by Eq. 11c11 The analytical solution of the level responses can be obtained from the inverse Laplace transform of Eq. 4 as: H (t ) =

Qi 1 t exp sin( t ) 2 H A

for < 1 where =

1 2 H

1 2

Qi t exp sinh( t ) 2 H A t Qi 1 exp = sinh( t ) 2 H A =

for > 1 where =

1 2 H

2 1

for = 1

(A11)

The peak of H(t) can be found in terms of â?¨ and â?śH : Qi H peak = g( ) (A12) A H

INSTRUMENTATION where g(␨) is:

1 x 2 + 1 x tan 1 x exp tan x g '( ) = 2 x 3 x

(

tan 1 x exp g( ) = 2 x 1+ x 2

)

1 2 1 x 2 1 (x tanh 1 x) exp tanh x = 2 x 3 x

1 2 for 0 < < 1 where x = = 2 exp( 1) for = 1 tanh 1 x 2 exp = x 1 x 2

for 0 < < 1 where x =

for > 1 where x =

2 1 for > 1 where x =

2 1

(A15)

(A13)

Finally, the constraint given by Eq. 11c can be expressed as: Qi g( ) H max (A14) A H Furthermore, g '( ) , which denotes dg(␨)/d␨ is given by:

Joonho Shin is a process systems engineer in corporate R&D at LG Chem Ltd., Korea. He holds a BS degree in chemical engineering from Korea University, and MS and PhD degrees in chemical engineering from KAIST. Dr. Shin began his professional career as a design and control specialist at SK Engineering & Construction in 1997. His industrial experience has focused on modeling, optimization and control of chemical and petrochemical industrial plants since 2002.

Moonyong Lee is a professor in the school of chemical engineering and technology at Yeungnam University, Korea. He holds a BS degree in chemical engineering from Seoul National University, and MS and PhD degrees in chemical engineering from KAIST. Dr. Lee worked in the refining and petrochemical plants of SK company for 10 years as a design and control specialist. Since joining the university in 1994, his areas of specialization have included modeling, design and chemical process control.

Jong Ku Lee is a vice president of LG Chem Research Park, LG Chem, Ltd., Korea. He holds a BS degree from Seoul National University, and MS and PhD degrees in chemical engineering from KAIST. Since joining LG Chem, Ltd in 1994, Dr. Lee has been working on numerous process modeling and optimization related projects. Currently he is in charge of the process modeling & solution group and doing research in the areas of energy savings and sustainability in chemical plants.

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Johnson Screens Europe . . . . . . . . .43

RS#

ONS . . . . . . . . . . . . . . . . . . . . . 26-27 (104)

Gulf Events . . . . . . . . . . . . . . . . . .85 (168)

Hunter Buildings . . . . . . . . . . . . . . .36 (154)

www.info.hotims.com/29416-75

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Ohmart/Vega . . . . . . . . . . . . . . . . .64 (163)

Gulf Books . . . . . . . . . . . . . . . . . .59 (170)

Haldor Topsoe A/S . . . . . . . . . . . . . .54

Company Website

Newton's . . . . . . . . . . . . . . . . . . . .45 (159)

Construction Boxscore . . . . . . . . . .57 (153)

www.info.hotims.com/29416-170

Aggreko . . . . . . . . . . . . . . . . . . . . .79 (167)

BASF Catalysts LLC . . . . . . . . . . . . .18

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www.info.hotims.com/29416-153

ABV Energy SpA . . . . . . . . . . . . . . .37 (155)

Axens . . . . . . . . . . . . . . . . . . . . . . .92

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Gulf Publishing Company

www.info.hotims.com/29416-69

Altair Strickland. . . . . . . . . . . . . . . .77

Company Website

(79)

Wood Group Surface Pumps . . . . . .39 (157) www.info.hotims.com/29416-157

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HPIN CONTROL ALLAN KERN, GUEST COLUMNIST kernag@yahoo.com

Back to the future: Process control in 2010 Where will process control money be spent in 2010? The choices include field instruments, safety systems (SISs), control system upgrades (DCSs), multivariable predictive control (MPC) and decision-support system applications (DSSs). In all cases, process control money is hard to spend. Itâ&#x20AC;&#x2122;s engineering-intensive, benefits can be elusive, and exactly which layer is most pressing, or most promising, is often in debate. As we enter a new decade, itâ&#x20AC;&#x2122;s a good time for a fresh look at this puzzle, because there are lessons to be learned from the past. And, for those seeking to leverage maximum overall improvement, as opposed to promoting a particular automation layer, there are clear choices. We have learned that decision-support systems require close coupling between IT and operations, not IT and automation. Weâ&#x20AC;&#x2122;ve learned that high MPC â&#x20AC;&#x153;service factorsâ&#x20AC;? donâ&#x20AC;&#x2122;t necessarily mean high utilization. And weâ&#x20AC;&#x2122;ve learned that upgrading DCS hardware is one thing, while capturing improvements such as effective operator graphics, best practice regulatory controls, and rational alarm configuration, is quite another. After two decades of activity, such insights are plentiful, but the overall message has been slow to take shape. For the past generation, MPC and DSSs have been the perceived frontier in process control progress. Meanwhile â&#x20AC;&#x153;base-layerâ&#x20AC;? health (field, SIS and DCS) has been considered settled territoryâ&#x20AC;&#x201D;something we now only â&#x20AC;&#x153;verifyâ&#x20AC;?, rather than something still demanding serious attention. But the message today is essentially the opposite. MPC and DSS have proven to be entirely tractable technologies, residing (as they do) fully in the software domain. Meanwhile, base-layer issues continue to rear their heads as the origin of most process control shortfalls, MPC and otherwise. Increasingly, the base layer is recognized as an area of renewed necessity, as well as opportunity. To make the shift in focus back to the base layer, it helps to reconsider its nature, which is not a simple matter of transmitters, valves and tuning, but a complex intersection of several intractable real-world inputs, including unanticipated process behaviors, process equipment that often requires careful kid-glove handling, and over-simple (or in the case of MPC, over-complex) control system configurations that too often fail to adequately address actual operating needs. Also, there is an offline, but nonetheless ongoing and real-time, parallel track of operational decision-making, best typified by the â&#x20AC;&#x153;morning meetingâ&#x20AC;?. When this activity is left unintegrated with automation plans (as it rather universally is), it has a pronounced tendency to defeat them at several points down the chain of command. It also helps to realize the great potential that base-layer controls have to bring about fundamental tangible improvement. Trip and incident reports continue to routinely cite shortcomings in operator graphics, alarm systems and regulatory controls as primary contributors. And many processes remain riddled with obvious opportunities in energy, yield and quality, despite a generation of MPC. Somewhat inexplicably, people often donâ&#x20AC;&#x2122;t look to their base-layer controls to address these problems, apparently because, 90

I FEBRUARY 2010 HYDROCARBON PROCESSING

DSS

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MPC

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Conventional automation â&#x20AC;&#x153;layers.â&#x20AC;?

over the past generation, updated DCS hardware with cookbook MPC have been considered the pinnacle of process control. With these in hand, the control system must be delivering all it can be expected to, they seem to be thinking. But this completely â&#x20AC;&#x153;misunderestimatesâ&#x20AC;?1 the necessities and the possibilities of the base layer. It also misses a big point about MPCâ&#x20AC;&#x201D;that it depends on a stable base layer, it does not provide one. The good news, especially for those experiencing lean times, is that a renewed base-layer strategy is not expensive. Some field work, as always, is indicated, but primarily the job is one of crafting better DCS (and in some cases SIS) software configurations. Those with a reasonably up-to-date DCS platform can hit the ground running in 2010 simply by allocating manpower, not money. But beware, trying to short-circuit the issue with a new contract or software application will probably leave you stuck in the â&#x20AC;&#x2122;90s â&#x20AC;&#x201C; long on budget, software and contractors, but (ironically) short-changed on base-layer performance. â&#x20AC;&#x153;Performance monitoringâ&#x20AC;? software will not sort this out for you, not this decade. The neo-control engineer, in this 2010 scenario, works with operations to devise high value-added automation solutionsâ&#x20AC;&#x201D;some unique or complex, others elementaryâ&#x20AC;&#x201D;that in most cases are appropriately deployed on the existing DCS platform, given the knowhow. A willingness to improve operation one bad-actor loop at a time when necessary is helpful. The bad news is this type of control engineer has been an endangered species since the 1980s, but perhaps they will stage a comeback in 2010. HP 1

LITERATURE CITED To misunderstand, causing to underestimate. President G. W. Bush, ca. 2001.

The author has 30 years of process control experience, including over a decade as MPC group leader at a major Middle East refinery, and has authored numerous articles on process control effectiveness. He is a professional engineer (inactive), a graduate of the University of Wyoming and a senior member of ISA.

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